U.S. patent number 9,512,047 [Application Number 15/076,512] was granted by the patent office on 2016-12-06 for ethylene-to-liquids systems and methods.
This patent grant is currently assigned to SILURIA TECHNOLOGIES, INC.. The grantee listed for this patent is Siluria Technologies, Inc.. Invention is credited to Joel Cizeron, Peter Czerpak, Carlos Faz, Jarod McCormick, William Michalak, Greg Nyce, Bipinkumar Patel, Guido Radaelli, Tim A. Rappold, Ron Runnebaum, Erik C. Scher, Aihua Zhang.
United States Patent |
9,512,047 |
Nyce , et al. |
December 6, 2016 |
Ethylene-to-liquids systems and methods
Abstract
Integrated systems are provided for the production of higher
hydrocarbon compositions, for example liquid hydrocarbon
compositions, from methane using an oxidative coupling of methane
system to convert methane to ethylene, followed by conversion of
ethylene to selectable higher hydrocarbon products. Integrated
systems and processes are provided that process methane through to
these higher hydrocarbon products.
Inventors: |
Nyce; Greg (Pleasanton, CA),
Czerpak; Peter (San Francisco, CA), Faz; Carlos
(Hayward, CA), McCormick; Jarod (San Carlos, CA),
Michalak; William (Redwood City, CA), Patel; Bipinkumar
(Richmond, TX), Radaelli; Guido (South San Francisco,
CA), Rappold; Tim A. (San Francisco, CA), Runnebaum;
Ron (Sacramento, CA), Scher; Erik C. (San Francisco,
CA), Zhang; Aihua (Daly City, CA), Cizeron; Joel
(Redwood City, CA) |
Applicant: |
Name |
City |
State |
Country |
Type |
Siluria Technologies, Inc. |
San Francisco |
CA |
US |
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Assignee: |
SILURIA TECHNOLOGIES, INC. (San
Francisco, CA)
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Family
ID: |
53524313 |
Appl.
No.: |
15/076,512 |
Filed: |
March 21, 2016 |
Prior Publication Data
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Document
Identifier |
Publication Date |
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US 20160200643 A1 |
Jul 14, 2016 |
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Related U.S. Patent Documents
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Application
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Filing Date |
Patent Number |
Issue Date |
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14789917 |
Jul 1, 2015 |
9321702 |
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14591850 |
Jan 7, 2015 |
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61925200 |
Jan 8, 2014 |
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62010986 |
Jun 11, 2014 |
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62050729 |
Sep 15, 2014 |
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Current U.S.
Class: |
1/1 |
Current CPC
Class: |
C10G
9/00 (20130101); C10G 11/00 (20130101); B01J
29/40 (20130101); C07C 5/327 (20130101); C07C
5/05 (20130101); C07C 1/0435 (20130101); F25J
3/0257 (20130101); C10G 57/02 (20130101); C07C
2/58 (20130101); F25J 3/0219 (20130101); C07C
1/06 (20130101); C07C 5/09 (20130101); F25J
3/0247 (20130101); F25J 3/0238 (20130101); C07C
1/0425 (20130101); C10L 3/10 (20130101); B01J
19/24 (20130101); C10G 50/00 (20130101); C07C
2/12 (20130101); C07C 1/12 (20130101); B01J
19/245 (20130101); C07C 2/84 (20130101); F25J
3/0209 (20130101); F25J 3/0233 (20130101); C07C
4/02 (20130101); C07C 2/84 (20130101); C07C
11/04 (20130101); C07C 2529/18 (20130101); F25J
2270/12 (20130101); F25J 2200/02 (20130101); F25J
2245/02 (20130101); F25J 2260/60 (20130101); F25J
2205/04 (20130101); F25J 2290/80 (20130101); C10G
2400/20 (20130101); F25J 2270/06 (20130101); F25J
2200/74 (20130101); F25J 2210/12 (20130101); F25J
2260/20 (20130101); F25J 2270/60 (20130101); F25J
2210/04 (20130101); Y02P 20/582 (20151101); B01J
2219/24 (20130101); Y02P 30/20 (20151101); B01J
2219/00103 (20130101); C07C 2529/70 (20130101); C07C
2529/85 (20130101); C07C 2523/46 (20130101); B01J
2219/00074 (20130101); C07C 2529/40 (20130101); F25J
2270/04 (20130101); Y02P 20/52 (20151101); C07C
2523/755 (20130101); F25J 2230/08 (20130101); F25J
2210/02 (20130101); Y02P 30/40 (20151101); C07C
2529/08 (20130101); F25J 2270/90 (20130101); C10G
2400/30 (20130101); F25J 2240/02 (20130101) |
Current International
Class: |
C07C
1/00 (20060101); C10L 3/10 (20060101); C07C
2/84 (20060101); C07C 5/327 (20060101); C07C
5/09 (20060101); C07C 5/05 (20060101); C07C
4/02 (20060101); C10G 50/00 (20060101); C10G
57/02 (20060101); C07C 1/12 (20060101); B01J
19/24 (20060101); C07C 2/12 (20060101); C07C
27/00 (20060101); C07C 1/04 (20060101) |
Field of
Search: |
;585/324 ;518/700 |
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Primary Examiner: Parsa; Jafar
Attorney, Agent or Firm: Wilson Sonsini Goodrich &
Rosati
Parent Case Text
CROSS-REFERENCE
This application is a continuation application of U.S. patent
application Ser. No. 14/789,917, filed Jul. 1, 2015, which
application is a continuation application of U.S. patent
application Ser. No. 14/591,850, filed Jan. 7, 2015, which
application claims the benefit of U.S. Provisional Patent
Application Ser. No. 61/925,200, filed Jan. 8, 2014, U.S.
Provisional Patent Application Ser. No. 62/010,986, filed Jun. 11,
2014, and U.S. Provisional Patent Application Ser. No. 62/050,729,
filed Sep. 15, 2014, each of which is entirely incorporated herein
by reference in its entirety.
Claims
What is claimed is:
1. A method for performing oxidative coupling of methane (OCM),
comprising: (a) directing oxygen (O.sub.2) and methane (CH.sub.4)
to an OCM unit that is upstream of a separations unit, wherein said
OCM unit is integrated with and in fluid communication with a
cracking unit, and wherein in said OCM unit, said O.sub.2 and said
CH.sub.4 react in an OCM process to yield compounds with two or
more carbon atoms (C.sub.2+ compounds), including ethylene
(C.sub.2H.sub.4); (b) directing at least a portion of said C.sub.2+
compounds to said cracking unit to yield a first stream comprising
(i) compounds having triple bonds or (ii) compounds having more
than one double bond, wherein said cracking unit operates
substantially adiabatically, and wherein said at least said portion
of said C.sub.2+compounds is directed to said cracking unit without
passage through said separations unit; and (c) directing said first
stream to said separations unit to yield a second stream having a
lower concentration of said (i) compounds having triple bonds or
(ii) compounds having more than one double bond with respect to
said first stream.
2. The method of claim 1, wherein said OCM unit has an inlet
temperature between about 450.degree. C. and about 600.degree.
C.
3. The method of claim 1, wherein said OCM unit has a pressure
between about 15 pounds per square inch gauge (psig) and about 125
psig.
4. The method of claim 1, wherein said OCM process has a C.sub.2+
selectivity of at least about 50%.
5. The method of claim 1, further comprising directing said second
stream to an ethylene-to-liquids (ETL) unit to convert
C.sub.2H.sub.4 in said second stream to yield an ETL product stream
comprising higher hydrocarbon product(s).
6. The method of claim 5, wherein said second stream, when directed
to said ETL unit, has less than about 100 parts per million (ppm)
of acetylene.
7. The method of claim 5, wherein said second stream, when directed
to said ETL unit, has less than or equal to about 5 mol % carbon
monoxide (CO).
8. The method of claim 5, wherein said ETL product stream has less
than about 50 weight percent (wt %) water.
9. The method of claim 5, wherein said ETL unit is operated at a
temperature greater than or equal to about 200.degree. C.
10. The method of claim 5, wherein said ETL unit is operated at a
pressure greater than or equal to about 200 pounds per square
inch.
11. The method of claim 5, further comprising directing at least a
portion of said ETL product stream to an additional separations
unit that recovers a liquid stream comprising said higher
hydrocarbon product(s) and a gas stream comprising hydrogen
(H.sub.2) and CO or carbon dioxide (CO.sub.2).
12. The method of claim 11, further comprising directing said gas
stream to a methanation unit to react said H.sub.2 with said CO or
CO.sub.2 to form CH.sub.4.
13. The method of claim 12, further comprising directing at least a
portion of said CH.sub.4 from said methanation unit to said OCM
unit.
14. The method of claim 1, wherein said at least said portion of
said C.sub.2+ compounds is directed to a treatment unit prior to
being directed to said cracking unit.
15. The method of claim 1, wherein an inlet temperature of said OCM
unit is at most about 600.degree. C.
16. The method of claim 1, wherein said OCM unit comprises a
plurality of OCM reactors.
17. The method of claim 1, wherein said cracking unit comprises a
plurality of cracking vessels.
18. The method of claim 1, further comprising directing one or more
alkanes to said cracking unit along a stream external to said OCM
unit.
19. The method of claim 1, wherein said first stream comprises said
compounds having triple bonds and said compounds having more than
one double bond.
Description
BACKGROUND
The modern petrochemical industry makes extensive use of cracking
and fractionation technology to produce and separate various
desirable compounds from crude oil. Cracking and fractionation
operations are energy intensive and generate considerable
quantities of greenhouse gases.
The gradual depletion of worldwide petroleum reserves and the
commensurate increase in petroleum prices may place extraordinary
pressure on refiners to minimize losses and improve efficiency when
producing products from existing feedstocks, and also to seek
viable alternative feedstocks capable of providing affordable
hydrocarbon intermediates and liquid fuels to downstream
consumers.
Methane may provide an attractive alternative feedstock for the
production of hydrocarbon intermediates and liquid fuels due to its
widespread availability and relatively low cost when compared to
crude oil. Worldwide methane reserves may be in the hundreds of
years at current consumption rates and new production stimulation
technologies may make formerly unattractive methane deposits
commercially viable.
Ethylene is an important commodity chemical intermediate. It may be
used in the production of polyethylene plastics, polyvinyl
chloride, ethylene oxide, ethylene chloride, ethylbenzene,
alpha-olefins, linear alcohols, vinyl acetate, and fuel blendstocks
such as, but not limited to, aromatics, alkanes and alkenes. With
economic growth in developed and developing portions of the world,
demand for ethylene and ethylene based derivatives continues to
increase. Currently, ethylene is produced through the cracking of
ethane derived either from crude oil distillates, called naphtha,
or from the relatively minor ethane component of natural gas.
Ethylene production is primarily limited to high volume production
as a commodity chemical in relatively large steam crackers or other
petrochemical complexes that also process the large number of other
hydrocarbon byproducts generated in the crude oil cracking process.
Producing ethylene from far more abundant and significantly less
expensive methane in natural gas provides an attractive alternative
to ethylene derived from ethane in natural gas or crude oil.
SUMMARY
Recognized herein is the need for efficient and commercially viable
systems and methods for converting ethylene to higher molecular
weight hydrocarbons, including gasoline, diesel fuel, jet fuel, and
aromatic chemicals. In some cases, the higher molecular weight
hydrocarbons can be produced from methane in an integrated process
that converts methane to ethylene and the ethylene to the higher
molecular weight compounds. An oxidative coupling of methane
("OCM") reaction is a process by which methane can form one or more
hydrocarbon compounds with two or more carbon atoms (also "C.sub.2+
compounds" herein), such as olefins like ethylene.
In an OCM process, methane can be oxidized to yield products
comprising C.sub.2+ compounds, including alkanes (e.g., ethane,
propane, butane, pentane, etc.) and alkenes (e.g., ethylene,
propylene, etc.). Such alkane (also "paraffin" herein) products may
not be suitable for use in downstream processes. Unsaturated
chemical compounds, such as alkenes (or olefins), may be more
preferable for use in downstream processes. Such compounds may be
polymerized to yield polymeric materials, which may be employed for
use in various commercial settings.
Oligomerization processes can be used to further convert ethylene
into longer chain hydrocarbons useful as polymer components for
plastics, vinyls, and other high value polymeric products.
Additionally, these oligomerization processes may be used to
convert ethylene to other longer hydrocarbons, such as C.sub.6,
C.sub.7, C.sub.8 and longer hydrocarbons useful for fuels like
gasoline, diesel, jet fuel and blendstocks for these fuels, as well
as other high value specialty chemicals.
An aspect of the present disclosure provides an oxidative coupling
of methane (OCM) system, comprising: (a) an OCM subsystem that (i)
takes as input a feed stream comprising methane (CH.sub.4) and a
feed stream comprising an oxidizing agent, and (ii) generates from
the methane and the oxidizing agent a product stream comprising
C.sub.2+ compounds and non-C.sub.2+ impurities; (b) at least one
separations subsystem downstream of, and fluidically coupled to,
the OCM subsystem, wherein the separations subsystem comprises a
first heat exchanger, a de-methanizer unit downstream of the first
heat exchanger, and a second heat exchanger downstream of the
de-methanizer unit, wherein (i) the first heat exchanger cools the
product stream, (ii) the de-methanizer unit accepts the product
stream from the first heat exchanger and generates an overhead
stream comprising at least a portion of the non-C.sub.2+
impurities, and (iii) at least a portion of the overhead stream is
cooled in the second heat exchanger and is subsequently directed to
the first heat exchanger to cool the product stream; and (c) an
olefin to liquids subsystem downstream of the OCM subsystem,
wherein the olefin to liquids subsystem is configured to generate
higher hydrocarbon(s) from one or more olefins included in the
C.sub.2+ compounds.
In some embodiments of aspects provided herein, the oxidizing agent
is O.sub.2. In some embodiments of aspects provided herein, the
O.sub.2 is provided by air. In some embodiments of aspects provided
herein, the OCM subsystem comprises at least one OCM reactor. In
some embodiments of aspects provided herein, the OCM subsystem
comprises at least one post-bed cracking unit downstream of the at
least one OCM reactor, which post-bed cracking unit is configured
to convert at least a portion of alkanes in the product stream to
alkenes. In some embodiments of aspects provided herein, the system
further comprises a non-OCM process upstream of the OCM subsystem.
In some embodiments of aspects provided herein, the non-OCM process
is a natural gas liquids process. In some embodiments of aspects
provided herein, the post-bed cracking unit is configured to
receive an additional stream comprising propane, separately from
the product stream. In some embodiments of aspects provided herein,
the non-C.sub.2+ impurities comprise one or more of nitrogen
(N.sub.2), oxygen (O.sub.2), water (H.sub.2O), argon (Ar), carbon
monoxide (CO), carbon dioxide (CO.sub.2), hydrogen (H.sub.2) and
methane (CH.sub.4). In some embodiments of aspects provided herein,
the higher hydrocarbon is a higher molecular weight
hydrocarbon.
An aspect of the present disclosure provides an oxidative coupling
of methane (OCM) system, comprising: (a) an OCM subsystem that (i)
takes as input a feed stream comprising methane (CH.sub.4) and a
feed stream comprising an oxidizing agent, and (ii) generates from
the methane and the oxidizing agent a product stream comprising
C.sub.2+ compounds and non-C.sub.2+ impurities; (b) at least one
methanation subsystem downstream of, and fluidically coupled to,
the OCM subsystem, wherein the methanation subsystem reacts CO,
CO.sub.2 and H.sub.2 included in the non-C.sub.2+ impurities to
generate methane; and (c) an ethylene-to-liquids (ETL) subsystem
downstream of the OCM subsystem, wherein the ETL subsystem is
configured to generate higher hydrocarbon(s) from ethylene included
in the C.sub.2+ compounds.
In some embodiments of aspects provided herein, at least a portion
of the methane generated in the methanation subsystem is recycled
to the OCM subsystem. In some embodiments of aspects provided
herein, the oxidizing agent is O.sub.2. In some embodiments of
aspects provided herein, the O.sub.2 is provided by air. In some
embodiments of aspects provided herein, the OCM subsystem comprises
at least one OCM reactor. In some embodiments of aspects provided
herein, the OCM subsystem comprises at least one post-bed cracking
unit downstream of the at least one OCM reactor, which post-bed
cracking unit is configured to convert at least a portion of
alkanes in the product stream to alkenes. In some embodiments of
aspects provided herein, the system further comprises a non-OCM
process upstream of the OCM subsystem. In some embodiments of
aspects provided herein, the non-OCM process is a natural gas
liquids process. In some embodiments of aspects provided herein,
the post-bed cracking unit is configured to receive an additional
stream comprising propane, separately from the product stream. In
some embodiments of aspects provided herein, the higher
hydrocarbon(s) comprise aromatics. In some embodiments of aspects
provided herein, the non-C.sub.2+ impurities comprise one or more
of nitrogen (N.sub.2), oxygen (O.sub.2), water (H.sub.2O), argon
(Ar), carbon monoxide (CO), carbon dioxide (CO.sub.2), hydrogen
(H.sub.2) and methane (CH.sub.4). In some embodiments of aspects
provided herein, the methanation subsystem comprises at least one
methanation reactor.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products, the method comprising: (a)
introducing methane and a source of oxidant into an oxidative
coupling of methane (OCM) reactor system capable of converting
methane to ethylene at reactor inlet temperatures of between about
450.degree. C. and 600.degree. C. and reactor pressures of between
about 15 psig and 125 psig, with C.sub.2+ selectivity of at least
about 50%, under conditions for the conversion of methane to
ethylene; (b) converting methane to a product gas comprising
ethylene; (c) introducing separate portions of the product gas into
at least a first and second integrated ethylene conversion reaction
systems, each integrated ethylene conversion reaction system being
configured for converting ethylene into a different higher
hydrocarbon product; and (d) converting the ethylene into different
higher hydrocarbon products.
In some embodiments of aspects provided herein, the first and
second integrated ethylene conversion systems are selected from the
group consisting of selective and full range ethylene conversion
systems. In some embodiments of aspects provided herein, the method
further comprises introducing a portion of the product gas into a
third integrated ethylene conversion system. In some embodiments of
aspects provided herein, the method further comprises introducing a
portion of the product gas into a fourth integrated ethylene
conversion systems. In some embodiments of aspects provided herein,
the at least first and second integrated ethylene conversion
systems are selected from the group consisting of linear alpha
olefin (LAO) systems, linear olefin systems, branched olefin
systems, saturated linear hydrocarbon systems, branched hydrocarbon
systems, saturated cyclic hydrocarbon systems, olefinic cyclic
hydrocarbon systems, aromatic hydrocarbon systems, oxygenated
hydrocarbon systems, halogenated hydrocarbon systems, alkylated
aromatic systems, and hydrocarbon polymer systems. In some
embodiments of aspects provided herein, the first and second
ethylene conversion systems are selected from the group consisting
of LAO systems that produce one or more of 1-butene, 1-hexene,
1-octene and 1-decene. In some embodiments of aspects provided
herein, at least one of the LAO systems is configured for
performing a selective LAO process. In some embodiments of aspects
provided herein, the first or second integrated ethylene conversion
systems comprises a full range ethylene oligomerization system
configured for producing higher hydrocarbons in the range of
C.sub.3 to C.sub.30. In some embodiments of aspects provided
herein, the OCM reactor system comprises nanowire OCM catalyst
material. In some embodiments of aspects provided herein, the
product gas comprises less than 5 mol % of ethylene. In some
embodiments of aspects provided herein, the product gas comprises
less than 3 mol % of ethylene. In some embodiments of aspects
provided herein, the product gas further comprises one or more
gases selected from the group consisting of CO.sub.2, CO, H.sub.2,
H.sub.2O, C.sub.2H.sub.6, CH.sub.4 and C.sub.3+ hydrocarbons. In
some embodiments of aspects provided herein, the method further
comprises enriching the product gas for ethylene prior to
introducing the separate portions of the product gas into the at
least first and second integrated ethylene conversion reaction
systems. In some embodiments of aspects provided herein, the method
further comprises introducing an effluent gas from the first or
second integrated ethylene conversion reaction systems into the OCM
reactor system.
An aspect of the present disclosure provides a method of producing
a plurality of liquid hydrocarbon products, the method comprising:
(a) catalytically converting methane to a product gas comprising
ethylene; and (b) processing separate portions of the product gas
with at least two discrete catalytic reaction systems selected from
the group consisting of linear alpha olefin (LAO) systems, linear
olefin systems, branched olefin systems, saturated linear
hydrocarbon systems, branched hydrocarbon systems, saturated cyclic
hydrocarbon systems, olefinic cyclic hydrocarbon systems, aromatic
hydrocarbon systems, oxygenated hydrocarbon systems, halogenated
hydrocarbon systems, alkylated aromatic systems, and hydrocarbon
polymer systems.
An aspect of the present disclosure provides a processing system,
comprising: (a) an oxidative coupling of methane (OCM) reactor
system comprising an OCM catalyst, the OCM reactor system being
fluidly connected at an input to a source of methane and a source
of oxidant, wherein the OCM reactor system (i) takes as input the
methane and the oxidant and (ii) generates from the methane and the
oxidant a product stream comprising C.sub.2+ compounds; (b) at
least a first catalytic ethylene conversion reactor systems and a
second catalytic ethylene conversion reactor system downstream of
the OCM reactor system, the first catalytic ethylene reactor system
being configured to convert ethylene to a first higher hydrocarbon,
and the second catalytic ethylene reactor system being configured
to convert ethylene to a second higher hydrocarbon different from
the first higher hydrocarbon; and (c) a selective coupling unit
between the OCM reactor system and the first and second catalytic
ethylene reactor systems, which selective coupling unit is
configured to selectively direct at least a portion of the product
gas to each of the first and second catalytic ethylene reactor
systems.
In some embodiments of aspects provided herein, the first and
second ethylene conversion systems are selected from the group
consisting of linear alpha olefin (LAO) systems, linear olefin
systems, branched olefin systems, saturated linear hydrocarbon
systems, branched hydrocarbon systems, saturated cyclic hydrocarbon
systems, olefinic cyclic hydrocarbon systems, aromatic hydrocarbon
systems, oxygenated hydrocarbon systems, halogenated hydrocarbon
systems, alkylated aromatic systems, ethylene copolymerization
systems, and hydrocarbon polymer systems. In some embodiments of
aspects provided herein, the OCM catalyst comprises a nanowire
catalyst. In some embodiments of aspects provided herein, the
system further comprises an ethylene recovery system between the
OCM reactor system and the first and second catalytic ethylene
conversion reactor systems, the ethylene recovery system configured
to enrich the product gas for ethylene.
An aspect of the present disclosure provides a chemical production
system, comprising: an OCM subsystem that includes an OCM reactor,
wherein the OCM reactor (i) takes as input a feed stream comprising
methane (CH.sub.4) and a feed stream comprising an oxidizing agent
and (ii) generates from the methane and the oxidizing agent
C.sub.2+ compounds and non-C.sub.2+ impurities; and an
ethylene-to-liquids (ETL) subsystem downstream of the OCM subsystem
that includes an ETL reactor, wherein the ETL reactor converts at
least a portion of the C.sub.2+ compounds to a product stream
comprising C.sub.3+ compounds, which C.sub.3+ compounds are
generated at a single pass conversion of at least about 40%.
In some embodiments of aspects provided herein, the methane is from
a non-OCM process. In some embodiments of aspects provided herein,
the ETL reactor operates at a pressure between about 4 bar and 50
bar. In some embodiments of aspects provided herein, the single
pass conversion is at least about 40% without recycle.
An aspect of the present disclosure provides a method for
generating hydrocarbons, comprising: (a) directing a feed stream
comprising methane (CH.sub.4) and a feed stream comprising an
oxidizing agent to an OCM reactor; (b) in the OCM reactor,
generating an OCM product stream comprising C.sub.2+ compounds and
non-C.sub.2+ impurities from the methane and the oxidizing agent;
(c) directing at least a portion of the C.sub.2+ compounds to an
ethylene-to-liquids (ETL) subsystem downstream of the OCM
subsystem, wherein the ETL subsystem has an ETL reactor that
converts at least a portion of the C.sub.2+ compounds in the OCM
product stream to an ETL product stream comprising C.sub.3+
compounds; and (d) recycling less than 25% of ethylene in the
product stream to the ETL subsystem.
In some embodiments of aspects provided herein, the OCM and ETL
subsystems generate the C.sub.3+ compounds at a single pass
conversion efficiency of at least about 40%. In some embodiments of
aspects provided herein, the single pass conversion efficiency is
at least about 40% without recycle. In some embodiments of aspects
provided herein, the methane is from a non-OCM process. In some
embodiments of aspects provided herein, the ETL reactor operates at
a pressure between about 10 bar and 50 bar.
An aspect of the present disclosure provides a method for
generating a catalyst, comprising: (a) providing a catalyst base
material having a first set of pores, wherein the base material
comprises an active component that facilitates the conversion of
olefins to a first set of hydrocarbons, at least some of which is
in liquid form at room temperature and atmospheric pressure; (b)
introducing a second set of pores into the base material having an
average diameter of at least about 10 nanometers as measured by BET
isotherms; and (c) providing one or more dopants on one or more
surfaces of the base material, wherein the one or more dopants
facilitate the conversion of olefins to a second set of
hydrocarbons, at least some of which are in liquid form at room
temperature and atmospheric pressure, wherein the second set of
hydrocarbons has a different product distribution than the first
set of hydrocarbons.
In some embodiments of aspects provided herein, the first set of
pores have an average diameter from at least about 4 Angstroms to
10 Angstroms. In some embodiments of aspects provided herein, the
base material comprises a zeolite. In some embodiments of aspects
provided herein, (b) is subsequent to (c). In some embodiments of
aspects provided herein, (b) and (c) are performed simultaneously.
In some embodiments of aspects provided herein, the base material
has a surface area from about 100 m.sup.2/g to 1000 m.sup.2/g. In
some embodiments of aspects provided herein, (c) comprises
providing dopants selected from the group consisting of Ga, Zn, Al,
In, Ni, Mg, B and Ag. In some embodiments of aspects provided
herein, the catalyst base material is H--Al-ZSM-5, H--Ga-ZSM-5,
H--Fe-ZSM-5, H--B-ZSM-5, or any combination thereof. In some
embodiments of aspects provided herein, the second set of
hydrocarbons has a narrower product distribution than the first set
of hydrocarbons.
An aspect of the present disclosure provides a system for
generating hydrocarbons, comprising: an ethylene-to-liquids (ETL)
unit comprising one or more ETL reactors, wherein an individual ETL
reactor accepts ethylene from a non-ETL process and generates a
product stream comprising higher hydrocarbons through an
oligomerization process, wherein at least some of the higher
hydrocarbons are in liquid form at room temperature and atmospheric
pressure; and at least one separations unit downstream of, and
fluidically coupled to, the ETL unit, wherein the separations unit
separates the product stream into individual streams, each
comprising a subset of the higher hydrocarbons.
In some embodiments of aspects provided herein, the ETL reactor
comprises a catalyst having an active material and one or more
dopants on surfaces of the active material. In some embodiments of
aspects provided herein, the system further comprises an oxidative
coupling of methane (OCM) unit upstream of the ETL unit, wherein
the OCM unit comprises one or more OCM reactors, each of which (i)
takes as input a feed stream comprising methane (CH.sub.4) and a
feed stream comprising an oxidizing agent, (ii) generates from the
methane and the oxidizing agent C.sub.2+ compounds and non-C.sub.2+
impurities, and (iii) directs at least a subset of ethylene in the
C.sub.2+ compounds to the ETL unit.
An aspect of the present disclosure provides a catalyst for the
conversion of ethylene to liquid hydrocarbon fuels, the catalyst
comprising: (a) a ZSM-5 base material; (b) a binder material; and
(c) a dopant material; wherein the catalyst has a cycle lifetime of
at least about 1 week when in contact with up to about 100 parts
per million (ppm) acetylene, and wherein the catalyst has a
replacement lifetime of at least about 1 year when in contact with
up to about 100 ppm acetylene.
An aspect of the present disclosure provides a catalyst for
hydrogenation of acetylene in an oxidative coupling of methane
(OCM) and ethylene to liquids (ETL) process comprising at least one
elemental metal, wherein the catalyst is capable of decreasing the
concentration of acetylene to less than about 100 parts per million
(ppm) in an OCM effluent prior to flowing the OCM effluent into an
ETL process.
In some embodiments of aspects provided herein, the catalyst is
capable of decreasing the concentration of acetylene to less than
about 10 ppm in the OCM effluent. In some embodiments of aspects
provided herein, the catalyst is capable of decreasing the
concentration of acetylene to less than about 1 ppm in the OCM
effluent. In some embodiments of aspects provided herein, the at
least one elemental metal includes palladium. In some embodiments
of aspects provided herein, the at least one elemental metal is
part of a metal oxide. In some embodiments of aspects provided
herein, the catalyst is capable of providing an OCM effluent that
comprises at least about 0.5% carbon monoxide. In some embodiments
of aspects provided herein, the catalyst is capable of providing an
OCM effluent that comprises at least about 1% carbon monoxide. In
some embodiments of aspects provided herein, the catalyst is
capable of providing an OCM effluent that comprises at least about
3% carbon monoxide. In some embodiments of aspects provided herein,
the catalyst has a lifetime of at least about 1 year. In some
embodiments of aspects provided herein, the catalyst is capable of
providing an OCM effluent that comprises at least about 0.1%
acetylene. In some embodiments of aspects provided herein, the
catalyst is capable of providing an OCM effluent that comprises at
least about 0.3% acetylene. In some embodiments of aspects provided
herein, the catalyst is capable of providing an OCM effluent that
comprises at least about 0.5% acetylene. In some embodiments of
aspects provided herein, the ETL process converts ethylene in the
OCM effluent into higher hydrocarbon(s). In some embodiments of
aspects provided herein, the at least one metal comprises a
plurality of metals.
An aspect of the present disclosure provides a catalyst for
converting carbon monoxide (CO) and/or carbon dioxide (CO.sub.2)
into methane (CH.sub.4) in an oxidative coupling of methane (OCM)
and ethylene to liquids (ETL) process, wherein the catalyst
comprises at least one elemental metal, and wherein the catalyst
converts CO and/or CO.sub.2 into CH.sub.4 at a selectivity for the
formation of methane that is at least about 10-fold greater than
the selectivity of the catalyst for formation of coke in an ETL
effluent.
In some embodiments of aspects provided herein, the catalyst has a
selectivity for the formation of methane that is at least about
100-fold greater than the selectivity of the catalyst for formation
of coke. In some embodiments of aspects provided herein, the
catalyst has a selectivity for the formation of methane that is at
least about 1000-fold greater than the selectivity of the catalyst
for formation of coke. In some embodiments of aspects provided
herein, the catalyst has a selectivity for the formation of methane
that is at least about 10000-fold greater than the selectivity of
the catalyst for formation of coke. In some embodiments of aspects
provided herein, the ETL effluent comprises at least about 3%
olefin and/or acetylene compounds. In some embodiments of aspects
provided herein, the ETL effluent comprises at least about 5%
olefin and/or acetylene compounds. In some embodiments of aspects
provided herein, the ETL effluent comprises at least about 10%
olefin and/or acetylene compounds. In some embodiments of aspects
provided herein, the at least one elemental metal includes nickel.
In some embodiments of aspects provided herein, the at least one
elemental metal is part of a metal oxide.
An aspect of the present disclosure provides a method for
preventing coke formation on a methanation catalyst in an oxidative
coupling of methane (OCM) and ethylene to liquids (ETL) process,
the method comprising: (a) providing an ETL effluent comprising
carbon monoxide (CO) and/or carbon dioxide (CO.sub.2); and (b)
using a methanation catalyst to perform a methanation reaction with
the ETL effluent, wherein: (i) hydrogen and/or water is added to
the ETL effluent prior to (b); (ii) olefins and/or acetylene in the
ETL effluent is hydrogenated prior to (b); and/or (iii) olefins
and/or acetylene are separated and/or condensed from the ETL
effluent prior to (b).
In some embodiments of aspects provided herein, (iii) is performed
using absorption or adsorption. In some embodiments of aspects
provided herein, the methanation reaction forms at least about
1000-fold more methane than coke. In some embodiments of aspects
provided herein, the methanation reaction forms at least about
10000-fold more methane than coke. In some embodiments of aspects
provided herein, the methanation reaction forms at least about
100000-fold more methane than coke. In some embodiments of aspects
provided herein, the method further comprises any two of (i), (ii)
and (iii). In some embodiments of aspects provided herein, the
method further comprises all of (i), (ii) and (iii). In some
embodiments of aspects provided herein, C.sub.5+ compounds are
removed from the ETL effluent prior to performing the methanation
reaction with the methanation catalyst. In some embodiments of
aspects provided herein, C.sub.4+ compounds are removed from the
ETL effluent prior to performing the methanation reaction with the
methanation catalyst. In some embodiments of aspects provided
herein, C.sub.3+ compounds are removed from the ETL effluent prior
to performing the methanation reaction with the methanation
catalyst.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products, comprising: (a) in an
oxidative coupling of methane (OCM) reactor, reacting methane and
an oxidant in an OCM process to yield heat and an OCM product
stream comprising hydrocarbon compounds with two or more carbon
atoms (C.sub.2+ compounds), including ethylene; (b) directing the
OCM product stream from the OCM reactor to a post-bed cracking
(PBC) unit downstream of the OCM reactor; (c) in the PBC unit,
subjecting the OCM product stream to thermal cracking under
conditions that are suitable to crack ethane to ethylene, wherein
the thermal cracking is conducted at least in part with the heat
from (a), thereby producing a PBC product stream comprising
ethylene and hydrogen (H.sub.2) at concentrations that are
increased relative to the respective concentrations of ethylene and
H.sub.2 in the OCM product stream; (d) directing the PBC product
stream from the PBC unit to an ethylene-to-liquids (ETL) reactor
downstream of the PBC unit, wherein the ETL reactor converts the
ethylene in the PBC product stream into higher hydrocarbons.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products, comprising: directing ethylene
and hydrogen (H.sub.2) into an ethylene-to-liquids (ETL) reactor,
wherein the ETL reactor is configured to convert hydrocarbon
compounds with two or more carbon atoms (C.sub.2+ compounds),
including ethylene, into higher hydrocarbons; and in the ETL
reactor, converting the ethylene into higher hydrocarbons in the
presence of the H.sub.2, wherein the converting results in less
coke formation than if the converting is conducted in the absence
of the H.sub.2.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products, comprising: directing ethylene
and water (H.sub.2O) into an ethylene-to-liquids (ETL) reactor,
wherein the ETL reactor is configured to convert hydrocarbon
compounds with two or more carbon atoms (C.sub.2+ compounds),
including ethylene, into higher hydrocarbons; and in an
ethylene-to-liquids (ETL) reactor, converting the ethylene into
higher hydrocarbons in the presence of the H.sub.2O, wherein the
converting results in less coke formation than if the converting is
conducted in the absence of the H.sub.2O.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products, comprising: (a) introducing a
feed stream comprising ethylene and ethane into an
ethylene-to-liquids (ETL) reactor, wherein the ETL reactor is
configured to convert hydrocarbon compounds with two or more carbon
atoms (C.sub.2+ compounds) into higher hydrocarbons, and wherein
the ethylene to ethane molar ratio in the feed stream is at least
about 3:1 and (b) in the ETL reactor, converting the ethylene into
the higher hydrocarbons.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products, comprising: directing ethylene
to an ethylene-to-liquids (ETL) reactor, wherein the ETL reactor is
configured to convert hydrocarbon compounds with two or more carbon
atoms (C.sub.2+ compounds) into higher hydrocarbons; in the ETL
reactor, converting the ethylene into the higher hydrocarbons; and
separating the higher hydrocarbons into at least two product
streams, at least one of which product streams is characterized by
five or more characteristics selected from the group consisting of:
(a) no more than 1.30 vol % benzene; (b) no more than 50 vol %
aromatics; (c) no more than 25 vol % olefins; (d) a motor octane
number (MON) of at least 82; (e) a total octane number of at least
87; (f) a Reid vapor pressure (RVP) of no more than 15 psi; (g) a
10% boiling point of no more than 70.degree. C.; (h) a 50% boiling
point of no more than 121.degree. C.; (i) a 90% boiling point of no
more than 190.degree. C.; (j) a final boiling point (FBP) of no
more than 221.degree. C.; and (k) an oxidative induction time of at
least 240 minutes.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products, comprising: directing ethylene
into an ethylene-to-liquids (ETL) reactor, wherein the ETL reactor
is configured to convert hydrocarbon compounds with two or more
carbon atoms (C.sub.2+ compounds), including ethylene, into higher
hydrocarbons; and in the ETL reactor, converting ethylene into
higher hydrocarbon products in an ETL product stream that comprises
less than 60% water.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products, comprising: (a) directing
ethylene into an ethylene-to-liquids (ETL) reactor, wherein the ETL
reactor comprises an ETL catalyst that is configured to convert
hydrocarbon compounds with two or more carbon atoms (C.sub.2+
compounds), including ethylene, into higher hydrocarbons; (b) in
the ETL reactor, converting the ethylene into higher hydrocarbons
to provide an ETL product stream comprising the higher
hydrocarbons, and forming coke on the ETL catalyst; (c) contacting
the ETL catalyst with an oxidant to regenerate the ETL catalyst by
burning the coke on the ETL catalyst; and (d) repeating (b)-(c) for
at least 20 cycles, wherein a composition of the ETL product stream
from a first cycle differs from a composition of the ETL product
stream from a twentieth cycle by no more than 0.1%.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products, comprising: (a) introducing a
feed stream comprising hydrocarbons into a fluid catalytic cracking
(FCC) reactor comprising an FCC catalyst, wherein the FCC catalyst
is configured to crack the hydrocarbons into lower molecular weight
hydrocarbons; (b) in the FCC reactor, (i) cracking the hydrocarbons
into the lower molecular weight hydrocarbons and (ii) generating
coke on the FCC catalyst; (c) transferring at least a portion of
the FCC catalyst into a regeneration unit and introducing an
oxidant stream into the regeneration unit; (d) in the regeneration
unit, burning the coke on the FCC catalyst in the presence of the
oxidant stream, thereby regenerating the FCC catalyst and producing
a flue gas stream comprising carbon monoxide and/or carbon dioxide;
(e) directing the flue gas stream into a heat exchanger to transfer
heat from the flue gas stream to a first stream comprising ethane
or propane; and (f) subjecting the first stream to thermal cracking
under conditions that (i) crack the ethane to ethylene and/or (ii)
crack the propane to propene, wherein the thermal cracking is
conducted at least in part with the heat from (e).
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products, comprising: (a) in a first
oxidative coupling of methane (OCM) reactor, reacting methane and a
first oxidant in an OCM process to yield a first OCM product stream
comprising unreacted methane and hydrocarbon compounds with two or
more carbon atoms (C.sub.2+ compounds), including ethylene; (b)
introducing the first OCM product stream into an
ethylene-to-liquids (ETL) reactor that is configured to convert
C.sub.2+ compounds into higher hydrocarbons; (c) in the ETL
reactor, converting at least a portion of the ethylene in the first
OCM product stream into higher hydrocarbons to provide an ETL
product stream comprising the higher hydrocarbons and the unreacted
methane; (d) introducing a second oxidant stream and at least a
portion of the ETL product stream into a second OCM reactor; and
(e) in the second OCM reactor, reacting the unreacted methane and
the second oxidant in another OCM process to yield a second OCM
product stream comprising C.sub.2+ compounds, including
ethylene.
In some embodiments of aspects provided herein, the method further
comprises, prior to the introducing of (e), removing at least a
portion of the higher hydrocarbons from the ETL product stream.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products, comprising: (a) directing a
feed stream comprising ethylene to an ethylene-to-liquids (ETL)
reactor, wherein the ETL reactor is configured to convert
hydrocarbon compounds with two or more carbon atoms (C.sub.2+
compounds) into higher hydrocarbons; (b) converting the ethylene to
an ETL product stream comprising the higher hydrocarbons; (c)
directing the ETL product stream to a separations system and, in
the separations system, separating the ETL product stream into a
higher hydrocarbon stream and a light olefin stream comprising
propylene and butene; (d) introducing the light olefin stream into
an oligomerization reactor, wherein the oligomerization reactor
includes an oligomerization catalyst that oligomerizes C.sub.2+
compounds into higher hydrocarbons; and (e) in the oligomerization
reactor, oligomerizing the propylene and butene in the light olefin
stream to produce an oligomerization product stream comprising
oligomerization products of propylene and butene.
In some embodiments of aspects provided herein, the oligomerization
product stream comprises olefins with carbon numbers from 6 to 16.
In some embodiments of aspects provided herein, a temperature
within the oligomerization reactor during the oligomerizing is from
about 50.degree. C. to 200.degree. C. In some embodiments of
aspects provided herein, the oligomerization catalyst comprises a
solid acid catalyst. In some embodiments of aspects provided
herein, the oligomerization reactor is of a form selected from the
group consisting of a slurry bed reactor, a fixed bed reactor, a
tubular isothermal reactor, a moving bed reactor, and a fluidized
bed reactor.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products, comprising: (a) in an
oxidative coupling of methane (OCM) reactor, reacting methane and
an oxidant in an OCM process to yield an OCM product stream
comprising unreacted methane and hydrocarbon compounds with two or
more carbon atoms (C.sub.2+ compounds) including ethylene, ethane,
and propane; (b) introducing the OCM product stream into an
ethylene-to-liquids (ETL) reactor, wherein the ETL reactor is
configured to convert the unreacted methane and at least a portion
of the C.sub.2+ compounds into aromatic hydrocarbons, and wherein
the ETL reactor comprises an ETL catalyst doped with one or more
dopants selected from the group consisting of molybdenum (Mo),
gallium (Ga), and tungsten (W); and (c) in the ETL reactor,
converting the unreacted methane and the at least the portion of
the C.sub.2+ compounds into an aromatic product stream comprising
the aromatic hydrocarbons.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products, comprising: (a) directing
hydrogen (H.sub.2) and a low octane stream comprising n-hexane to
an isomerization reactor that is configured to isomerize n-hexane
to i-hexane, wherein the low octane stream is characterized by an
octane number of no more than 62; and (b) reacting the H.sub.2 and
the n-hexane to produce an isomerization product stream comprising
i-hexane, wherein the isomerization product stream is characterized
by an octane number of at least 73.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products, comprising: (a) in a natural
gas liquids (NGL) system, producing from natural gas an NGL product
stream comprising hydrocarbon compounds with four or more carbon
atoms (C.sub.4+ compounds), including butanes; (b) introducing the
first NGL product stream into an isomerization reactor configured
to isomerize the C.sub.4+ compounds); and (c) in the isomerization
reactor, isomerizing at least a portion of the C.sub.4+ compounds
to form isomerization products, thereby producing an isomerate
stream comprising the isomerization products.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products, comprising: (a) in an
oxidative coupling of methane (OCM) reactor, reacting methane and
an oxidant in an OCM process to yield an OCM product stream
comprising unreacted methane and hydrocarbon compounds comprising
two or more carbon atoms (C.sub.2+ compounds), including ethylene;
(b) introducing the OCM product stream into an ethylene-to-liquids
(ETL) reactor that reacts the ethylene in the OCM product stream to
yield an ETL product stream higher hydrocarbons and unreacted
methane; and (c) introducing the ETL product stream into at least
one separation unit that separates the ETL product stream into a
gas stream comprising the unreacted methane and at least one
product stream comprising hydrocarbon compounds with at least 3, 4,
or 5 carbon atoms.
In some embodiments of aspects provided herein, the methane is
supplied at least in part from a natural gas pipeline, and wherein
the method further comprises outputting the gas stream to the
natural gas pipeline. In some embodiments of aspects provided
herein, the methane is supplied at least in part from a cryogenic
separations system, and wherein the method further comprises
directing the gas stream to a re-compressor unit. In some
embodiments of aspects provided herein, the methane is supplied at
least in part from a cryogenic separations system, and wherein the
method further comprises compressing the gas stream in a compressor
to produce a compressed stream and directing the compressed stream
to the cryogenic separations unit. In some embodiments of aspects
provided herein, the methane is supplied at least in part from a
cryogenic separations unit, and the method further comprises:
compressing the gas stream in a compressor to produce a compressed
stream; directing the compressed stream to the cryogenic
separations unit; in the cryogenic separations unit, removing any
C.sub.2+ compounds from the gas stream along a C.sub.2+ product
stream; and optionally directing the gas stream to a re-compressor
unit.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products including hydrocarbon compounds
with two carbon atoms (C.sub.2 compounds), hydrocarbon compounds
with three carbon atoms (C.sub.3 compounds), hydrocarbon compounds
with four carbon atoms (C.sub.4 compounds), and hydrocarbon
compounds with five or more carbon atoms (C.sub.5+ compounds),
comprising: (a) introducing a natural gas stream comprising methane
into a gas treatment system and, in the gas treatment system,
removing from the natural gas stream at least one of mercury,
water, and acid gases; (b) directing the natural gas stream from
the gas treatment system into a natural gas liquids (NGL)
extraction system that produces from the natural gas stream a first
stream comprising methane and a second stream comprising C.sub.2
compounds, C.sub.3 compounds, C.sub.4 compounds, and C.sub.5+
compounds; (c) directing a first portion of the first stream into a
liquefaction unit, and in the liquefaction unit, producing liquid
natural gas from the first portion of the first stream; (d)
directing the second stream into an NGL fractionation system that
separates the second stream into at least (i) a C.sub.2-C.sub.3
stream comprising C.sub.2 compounds and C.sub.3 compounds, (ii) a
C.sub.4 stream comprising C.sub.4 compounds, and (iii) a C.sub.5+
stream comprising C.sub.5+ compounds; (e) directing a second
portion of the first stream, the C.sub.2-C.sub.3 stream, and an
oxidant into an oxidative coupling of methane (OCM) system that
converts the methane in the second portion of the first stream in
an OCM process to yield an OCM product stream including ethylene;
(f) directing the OCM product stream into an ethylene-to-liquids
(ETL) reactor that converts the ethylene in the OCM product stream
into the higher hydrocarbons, thereby forming an ETL product stream
comprising C.sub.2 compounds, C.sub.3 compounds, C.sub.4 compounds,
and C.sub.5+ compounds; and (g) directing the ETL product stream
into the NGL extraction system.
In some embodiments of aspects provided herein, the method further
comprises, prior to the directing of (b), directing the natural gas
stream from the gas treatment system into a pre-cooling system,
and, in the pre-cooling system, removing a first fuel gas stream
comprising methane from the natural gas stream. In some embodiments
of aspects provided herein, the method further comprises directing
the liquid natural gas stream into a nitrogen rejection unit, and,
in the nitrogen rejection unit, removing a stream comprising
nitrogen from the liquid natural gas stream.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products including hydrocarbon compounds
with two carbon atoms (C.sub.2 compounds), hydrocarbon compounds
with three carbon atoms (C.sub.3 compounds), hydrocarbon compounds
with four carbon atoms (C.sub.4 compounds), and hydrocarbon
compounds with five or more carbon atoms (C.sub.5+ compounds),
comprising: (a) directing a natural gas stream into a natural gas
liquids (NGL) extraction system that produces from the natural gas
stream a first stream comprising methane and a second stream
comprising C.sub.2 compounds, C.sub.3 compounds, C.sub.4 compounds,
and C.sub.5+ compounds; (b) removing a first portion of the first
stream as a pipeline gas product stream; (c) directing the second
stream into an NGL fractionation system that separates the second
stream into at least (i) a C.sub.2-C.sub.3 stream comprising
C.sub.2 compounds and C.sub.3 compounds, (ii) a C.sub.4 stream
comprising C.sub.4 compounds, and (iii) a C.sub.5+ stream
comprising C.sub.5+ compounds; (d) directing a second portion of
the first stream, the C.sub.2-C.sub.3 stream, and an oxidant into
an oxidative coupling of methane (OCM) system that converts the
methane in the second portion of the first stream in an OCM process
to yield an OCM product stream including ethylene; (e) directing
the OCM product stream into an ethylene-to-liquids (ETL) reactor
that converts the ethylene in the OCM product stream into an ETL
product stream comprising C.sub.2 compounds, C.sub.3 compounds,
C.sub.4 compounds, and C.sub.5+ compounds; and (f) directing the
ETL product stream into the NGL extraction system.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products including hydrocarbon compounds
with two carbon atoms (C.sub.2 compounds"), hydrocarbon compounds
with three carbon atoms (C.sub.3 compounds), hydrocarbon compounds
with four carbon atoms (C.sub.4 compounds), and hydrocarbon
compounds with five or more carbon atoms (C.sub.5+ compounds),
comprising: (a) introducing a first natural gas stream comprising
methane into a gas treatment system that removes from the first
natural gas stream at least one of mercury, water, and acid gases;
(b) introducing a second natural gas stream comprising methane into
a gas conditioning system that removes from the second natural gas
stream at least one sulfur compound; (c) directing the first
natural gas stream from the gas treatment system and a first
portion of the second natural gas stream from the gas conditioning
system into a natural gas liquids (NGL) extraction system that
produces from the first natural gas stream and the first portion of
the second natural gas stream (i) a first stream comprising
methane, (ii) a second stream comprising C.sub.2 compounds, and
(iii) a third stream comprising C.sub.2 compounds, C.sub.3
compounds, C.sub.4 compounds, and C.sub.5+ compounds, wherein a
portion of the first stream is removed as a pipeline gas product
stream; (d) directing the third stream into an NGL fractionation
system that separates the third stream into at least (i) a C.sub.2
stream comprising C.sub.2 compounds, (ii) a C.sub.3-C.sub.4 stream
comprising C.sub.3 compounds and C.sub.4 compounds, and (iii) a
C.sub.5+ stream comprising C.sub.5+ compounds; (e) directing a
second portion of the second natural gas stream from the gas
conditioning system, the second stream from the NGL extraction
system, the C.sub.2 stream from the NGL fractionation system, and
an oxidant into an oxidative coupling of methane (OCM) reactor that
converts methane at least some of the streams in an OCM process to
yield an OCM product stream including ethylene; (f) directing the
OCM product stream into an ethylene-to-liquids (ETL) reactor that
converts the ethylene in the OCM product stream into an ETL product
stream comprising C.sub.2 compounds, C.sub.3 compounds, C.sub.4
compounds, and C.sub.5+ compounds; and (g) directing the ETL
product stream into the NGL extraction system.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products including hydrocarbon compounds
with two carbon atoms (C.sub.2 compounds), hydrocarbon compounds
with three carbon atoms (C.sub.3 compounds), hydrocarbon compounds
with four carbon atoms (C.sub.4 compounds), and hydrocarbon
compounds with five or more carbon atoms (C.sub.5+ compounds),
comprising: (a) directing a first stream including ethylene from a
refinery gas plant into an ethylene-to-liquids (ETL) reactor that
converts the ethylene into an ETL product stream comprising C.sub.2
compounds, C.sub.3 compounds, C.sub.4 compounds, and C.sub.5+
compounds; (b) directing the ETL product stream into a separations
system that separates the ETL product stream into at least (i) a
fuel gas stream comprising methane, (ii) a C.sub.2 stream
comprising C.sub.2 compounds, and (iii) a C.sub.3 stream comprising
C.sub.3 compounds; (c) using a heat exchanger, transferring heat
from the C.sub.2 stream to a first stream comprising ethane and/or
propane; and (d) subjecting the first stream to thermal cracking
under conditions that crack the ethane to ethylene and/or the
propane to propene, wherein the thermal cracking is conducted at
least in part with the heat from (c).
In some embodiments of aspects provided herein, the method further
comprises directing the C.sub.2 stream from the heat exchanger to
the ETL reactor.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products including hydrocarbon compounds
with two carbon atoms (C.sub.2 compounds), hydrocarbon compounds
with three carbon atoms (C.sub.3 compounds), hydrocarbon compounds
with four carbon atoms (C.sub.4 compounds), and hydrocarbon
compounds with five or more carbon atoms (C.sub.5+ compounds),
comprising: (a) directing a first stream including ethylene from a
refinery gas plant into a demethanizer that removes a first methane
stream including methane from the first stream, wherein the first
stream is subjected to sulfur removal prior to being directed to
the demethanizer; (b) directing the first methane stream, a second
methane stream comprising methane, and an oxidant into an oxidative
coupling of methane (OCM) system that converts methane in the in an
OCM process to yield an OCM product stream including ethylene; (c)
directing the first stream and the OCM product stream into an
ethylene-to-liquids (ETL) that converts the ethylene in the OCM
product stream into an ETL product stream comprising C.sub.2
compounds, C.sub.3 compounds, C.sub.4 compounds, and C.sub.5+
compounds; (d) directing the ETL product stream into a separations
system that separates the ETL product stream into at least streams
comprising a C.sub.2-C.sub.3 stream comprising C.sub.2 compounds
and C.sub.3 compounds; and (e) using a heat exchanger, transferring
heat from the C.sub.2-C.sub.3 stream to a second stream comprising
ethane and/or propane; and (f) subjecting the second stream to
thermal cracking under conditions that crack the ethane to ethylene
and/or the propane to propene, wherein the thermal cracking is
conducted at least in part with the heat from (e).
In some embodiments of aspects provided herein, the method further
comprises directing at least a portion of the C.sub.2-C.sub.3
stream from the heat exchanger to the OCM reactor system. In some
embodiments of aspects provided herein, the method further
comprises directing at least a portion of the C.sub.2-C.sub.3
stream from the heat exchanger to the ETL reactor.
An aspect of the present disclosure provides a method of producing
a plurality of hydrocarbon products including hydrocarbon compounds
with two or more carbon atoms (C.sub.2+ compounds), comprising: (a)
directing methane and an oxidant to an oxidative coupling of
methane (OCM) reactor that is upstream of a post-bed cracking (PBC)
unit, wherein the OCM reactor is configured to facilitate an OCM
reaction using the methane and the oxidant to generate the C.sub.2+
compounds including ethylene and one or more alkanes, and wherein
the PBC unit is configured to convert the one or more alkanes,
including ethane, to one or more alkenes, including ethylene; (b)
in the OCM reactor, reacting the methane and the oxidant in the OCM
reaction to generate an OCM product stream and heat, wherein the
OCM product stream comprises ethylene and one or more alkanes; (c)
directing the OCM product stream to the PBC unit; (d) in the PBC
unit, subjecting the OCM product stream to thermal cracking under
conditions that crack ethane to ethylene, wherein the thermal
cracking is conducted at least in part with the heat from (c),
thereby producing a PBC product stream comprising ethylene; (e)
directing the PBC product stream to a separations module, and, in
the separations module, separating ethane from the PBC product
stream to generate an ethane stream; and (f) directing the ethane
stream to the PBC unit.
Additional aspects and advantages of the present disclosure will
become readily apparent to those skilled in this art from the
following detailed description, wherein only illustrative
embodiments of the present disclosure are shown and described. As
will be realized, the present disclosure is capable of other and
different embodiments, and its several details are capable of
modifications in various obvious respects, all without departing
from the disclosure. Accordingly, the drawings and description are
to be regarded as illustrative in nature, and not as
restrictive.
INCORPORATION BY REFERENCE
All publications, patents, and patent applications mentioned in
this specification are herein incorporated by reference to the same
extent as if each individual publication, patent, or patent
application was specifically and individually indicated to be
incorporated by reference.
BRIEF DESCRIPTION OF THE DRAWINGS
The novel features of the invention are set forth with
particularity in the appended claims. A better understanding of the
features and advantages of the present invention will be obtained
by reference to the following detailed description that sets forth
illustrative embodiments, in which the principles of the invention
are utilized, and the accompanying drawings or figures (also "FIG."
and "FIGs." herein), of which:
FIG. 1 shows an oxidative coupling of methane (OCM) reactor
system;
FIG. 2 schematically illustrates differentially cooled tubular
reactor systems;
FIG. 3 schematically illustrates a reactor system with two or more
tubular reactors;
FIG. 4A schematically illustrates an alternative approach for
varying reactor volumes in order to vary residence time of
reactants in a catalyst bed;
FIG. 4B schematically illustrates an exemplary fluidized bed
reactor;
FIG. 4C schematically illustrates exemplary moving bed, fluidized
bed, and slurry bed reactors;
FIG. 5 is an example of the manner in which product distribution
can change over time for an ETL catalyst;
FIG. 6 schematically illustrates an ethylene-to-liquids (ETL)
reactor system with process inlet and recycle stream combining to
form a reactor inlet process stream;
FIG. 7A shows liquid phase hydrocarbon yield as a function of
C.sub.2H.sub.4 conversion by using a single pass reactor;
FIG. 7B shows liquid phase hydrocarbon yield as a function of
C.sub.2H.sub.4 conversion by using a reactor with 5:1 recycle:fresh
ratio;
FIG. 8 is a plot showing increasing C.sub.5+ yield (liquid
condensed at about 0.degree. C.) with increasing recycle reaction
conditions;
FIG. 9 shows an example of a pressure swing adsorption (PSA)
unit;
FIG. 10 schematically illustrates an integrated OCM system with
integrated separations system;
FIG. 11 shows an example of NGL extraction in a liquefied natural
gas (LNG) facility;
FIG. 12 shows an integrated OCM-ETL system for use in LNG
production;
FIG. 13 shows the system of FIG. 12 that has been modified for use
with a diluted C1 (methane) fuel gas stream;
FIG. 14 shows an example OCM-ETL system comprising OCM and ETL
sub-systems, and a separations sub-system downstream of the ETL
sub-system;
FIG. 15 shows an OCM-ETL system comprising OCM and ETL sub-systems,
and a cryogenic cold box downstream of the ETL sub-system;
FIG. 16 shows another OCM-ETL in an alternative configuration to
that shown in FIG. 14;
FIG. 17 show examples of OCM-ETL midstream integration;
FIG. 18 show examples of OCM-ETL midstream integration;
FIG. 19 shows OCM-ETL systems with various skimmer and recycle
configurations.
FIG. 20 shows an example of ETL integration in a refinery;
FIG. 21 shows another example of ETL integration in a refinery;
FIG. 22 shows another example of ETL integration in a refinery;
FIG. 23A schematically illustrates a natural gas liquids (NGL)
system; FIG. 23B schematically illustrates the NGL process of FIG.
23A retrofitted with an OCM and ethylene to liquids system;
FIG. 24 schematically illustrates an oxidative coupling of methane
(OCM) olefins to liquids process integrated in an NGL system,
employing air in an OCM process;
FIG. 25 schematically illustrates an OCM-ETL integration with an
existing NGL system, employing oxygen (O.sub.2) in an OCM
process;
FIG. 26 schematically illustrates a methanation system;
FIG. 27 shows an example of methanation systems for OCM and
ETL;
FIG. 28 shows a separation system that may be employed for use with
systems and methods of the present disclosure;
FIG. 29 shows another separation system that may be employed for
use with systems and methods of the present disclosure;
FIG. 30 shows another separation system that may be employed for
use with systems and methods of the present disclosure;
FIG. 31 shows another separation system that may be employed for
use with systems and methods of the present disclosure;
FIG. 32 shows an example of an ethane skimmer implementation of OCM
and ETL; and
FIG. 33 shows a computer system that is programmed or otherwise
configured to regulate OCM reactions;
FIG. 34 schematically illustrates a process flow for conversion of
ethylene to higher liquid hydrocarbons for use in, e.g., fuels and
fuel blendstocks;
FIG. 35 shows a graph of exemplary product compositions over time
on stream;
FIGS. 36A-36E shows graphs of ETL products with various feedstocks.
FIG. 36A shows a graph of ETL product with ethylene feedstock; FIG.
36B shows a graph of ETL product with propylene feedstock; FIG. 36C
shows a graph of ETL product with butylene feedstock; FIG. 36D
shows a graph of ETL product with 50:50 ethylene/propylene
feedstock; and FIG. 36E shows a graph of ETL product with 50:50
ethylene/butylene feedstock;
FIG. 37 shows a graph of product composition versus catalyst bed
peak temperature;
FIG. 38 shows a graph of ETL product divided into a gasoline
fraction and a jet fraction;
FIG. 39 shows a graph of crush strength for various catalyst
formulations;
FIG. 40 shows a graph comparing catalyst aging under commercial and
accelerated conditions;
FIG. 41 shows a graph comparing product composition over catalyst
regeneration cycles; and
FIG. 42 shows an OCM reactor comprising an integrated catalyst unit
and cracking unit.
DETAILED DESCRIPTION
While various embodiments of the invention have been shown and
described herein, it will be obvious to those skilled in the art
that such embodiments are provided by way of example only. Numerous
variations, changes, and substitutions may occur to those skilled
in the art without departing from the invention. It should be
understood that various alternatives to the embodiments of the
invention described herein may be employed.
Unless the context requires otherwise, throughout the specification
and claims which follow, the word "comprise" and variations
thereof, such as, "comprises" and "comprising" are to be construed
in an open, inclusive sense, that is, as "including, but not
limited to." Further, headings provided herein are for convenience
only and do not interpret the scope or meaning of the claimed
invention.
Reference throughout this specification to "one embodiment" or "an
embodiment" means that a particular feature, structure or
characteristic described in connection with the embodiment is
included in at least one embodiment. Thus, the appearances of the
phrases "in one embodiment" or "in an embodiment" in various places
throughout this specification are not necessarily all referring to
the same embodiment. Furthermore, the particular features,
structures, or characteristics may be combined in any suitable
manner in one or more embodiments. Also, as used in this
specification and the appended claims, the singular forms "a,"
"an," and "the" include plural referents unless the content clearly
dictates otherwise. It should also be noted that the term "or" is
generally employed in its sense including "and/or" unless the
content clearly dictates otherwise.
The term "OCM process," as used herein, generally refers to a
process that employs or substantially employs an oxidative coupling
of methane (OCM) reaction. An OCM reaction can include the
oxidation of methane to a higher hydrocarbon (e.g., higher
molecular weight hydrocarbon or higher chain hydrocarbon) and
water, and involves an exothermic reaction. In an OCM reaction,
methane can be partially oxidized to one or more C.sub.2+
compounds, such as ethylene, propylene, butylenes, etc. In an
example, an OCM reaction is
2CH.sub.4+O.sub.2.fwdarw.C.sub.2H.sub.4+2H.sub.2O. An OCM reaction
can yield C.sub.2+ compounds. An OCM reaction can be facilitated by
a catalyst, such as a heterogeneous catalyst. Additional
by-products of OCM reactions can include CO, CO.sub.2, H.sub.2, as
well as hydrocarbons, such as, for example, ethane, propane,
propene, butane, butene, and the like.
The term "non-OCM process," as used herein, generally refers to a
process that does not employ or substantially employ an oxidative
coupling of methane reaction. Examples of processes that may be
non-OCM processes include non-OCM hydrocarbon processes, such as,
for example, non-OCM processes employed in hydrocarbon processing
in oil refineries, a natural gas liquids separations processes,
steam cracking of ethane, steam cracking or naphtha,
Fischer-Tropsch processes, and the like.
The term "ethylene-to-liquids" (ETL), as used herein, generally
refers to any device, system, method (or process) that can convert
an olefin (e.g., ethylene) to higher molecular weight hydrocarbons,
which can be in liquid form.
The term "non-ETL process," as used herein, generally refers to a
process that does not employ or substantially employ the conversion
of an olefin to a higher molecular weight hydrocarbon through
oligomerization. Examples of processes that may be non-ETL
processes include processes employed in hydrocarbon processing in
oil refineries, a natural gas liquids separations processes, steam
cracking of ethane, steam cracking or naphtha, Fischer-Tropsch
processes, and the like.
The terms "C.sub.2+" and "C.sub.2+ compound," as used herein,
generally refer to a compound comprising two or more carbon atoms,
e.g., C.sub.2, C.sub.3 etc. C.sub.2+ compounds include, without
limitation, alkanes, alkenes, alkynes and aromatics containing two
or more carbon atoms. In some cases, C.sub.2+ compounds include
aldehydes, ketones, esters and carboxylic acids. Examples of
C.sub.2+ compounds include ethane, ethene, acetylene, propane,
propene, butane, butene, etc.
The term "non-C.sub.2+ impurities," as used herein, generally
refers to material that does not include C.sub.2+ compounds.
Examples of non-C.sub.2+ impurities, which may be found in certain
OCM reaction product streams, include nitrogen (N.sub.2), oxygen
(O.sub.2), water (H.sub.2O), argon (Ar), hydrogen (H.sub.2) carbon
monoxide (CO), carbon dioxide (CO.sub.2) and methane
(CH.sub.4).
The term "weight hourly space velocity" (WHSV), as used herein,
generally refers to the mass flow rate of olefins in a feed divided
by the mass of a catalyst, which can have units of inverse time
(e.g., hr.sup.-1).
The term "slate," as used herein, generally refers to distribution,
such as product distribution.
The term "oligomerization," as used herein, generally refers to a
reaction in which hydrocarbons are combined to form larger chain
hydrocarbons.
The term "greenfield," as used herein, generally refers to an
investment in a manufacturing, office, industrial or other physical
commerce-related structure or group of structures in an area where
no previous facilities exist or have existed.
The term "brownfield," as used herein, generally refers to an
investment at a site that was previously used for business
purposes, such as a steel mill or an oil refinery, but is
subsequently expanded or upgraded to achieve a return.
The term "catalyst," as used herein, generally refers to a
substance that alters the rate of a chemical reaction. A catalyst
may either increase the chemical reaction rate (i.e. a "positive
catalyst") or decrease the reaction rate (i.e. a "negative
catalyst"). A catalyst can be a heterogeneous catalyst. Catalysts
can participate in a reaction in a cyclic fashion such that the
catalyst is cyclically regenerated. "Catalytic" generally means
having the properties of a catalyst.
The term "nanowire," as used herein, generally refers to a nanowire
structure having at least one diameter on the order of nanometers
(e.g. between about 1 and 100 nanometers) and an aspect ratio
greater than 10:1. The "aspect ratio" of a nanowire is the ratio of
the actual length (L) of the nanowire to the diameter (D) of the
nanowire. Aspect ratio is expressed as L:D.
The term "polycrystalline nanowire," as used herein, generally
refers to a nanowire having multiple crystal domains.
Polycrystalline nanowires generally have different morphologies
(e.g. bent vs. straight) as compared to the corresponding
"single-crystalline" nanowires.
The term "effective length" of a nanowire, as used herein,
generally refers to the shortest distance between the two distal
ends of a nanowire as measured by transmission electron microscopy
(TEM) in bright field mode at 5 kilo electron volt (keV). "Average
effective length" refers to the average of the effective lengths of
individual nanowires within a plurality of nanowires.
The term "actual length" of a nanowire, as used herein, generally
refers to the distance between the two distal ends of a nanowire as
traced through the backbone of the nanowire as measured by TEM in
bright field mode at 5 keV. "Average actual length" refers to the
average of the actual lengths of individual nanowires within a
plurality of nanowires.
A "diameter" of a nanowire can be measured in an axis perpendicular
to the axis of the nanowire's actual length (i.e. perpendicular to
the nanowires backbone). The diameter of a nanowire will vary from
narrow to wide as measured at different points along the nanowire
backbone. As used herein, the diameter of a nanowire is the most
prevalent (i.e. the mode) diameter.
A "ratio of effective length to actual length" can be determined by
dividing the effective length by the actual length. A nanowire
having a "bent morphology" can have a ratio of effective length to
actual length of less than one as described in more detail herein.
A straight nanowire will have a ratio of effective length to actual
length equal to one.
The term "inorganic," as used herein, generally refers to a
substance comprising a metal element or semi-metal element. In
certain embodiments, inorganic refers to a substance comprising a
metal element. An inorganic compound can contain one or more metals
in its elemental state, or more typically, a compound formed by a
metal ion (M.sup.n+, wherein n 1, 2, 3, 4, 5, 6 or 7) and an anion
(X.sup.m-, m is 1, 2, 3 or 4), which balance and neutralize the
positive charges of the metal ion through electrostatic
interactions. Non-limiting examples of inorganic compounds include
oxides, hydroxides, halides, nitrates, sulfates, carbonates,
phosphates, acetates, oxalates, and combinations thereof, of metal
elements. Other non-limiting examples of inorganic compounds
include Li.sub.2CO.sub.3, Li.sub.2PO.sub.4, LiOH, Li.sub.2O, LiCl,
LiBr, LiI, Li.sub.2C.sub.2O.sub.4, Li.sub.2SO.sub.4,
Na.sub.2CO.sub.3, Na.sub.2PO.sub.4, NaOH, Na.sub.2O, NaCl, NaBr,
NaI, Na.sub.2C.sub.2O.sub.4, Na.sub.2SO.sub.4, K.sub.2CO.sub.3,
K.sub.2PO.sub.4, KOH, K.sub.2O, KCl, KBr, KI,
K.sub.2C.sub.2O.sub.4, K.sub.2SO.sub.4, Cs.sub.2CO.sub.3,
CsPO.sub.4, CsOH, Cs.sub.2O, CsCl, CsBr, CsI, CsC.sub.2O.sub.4,
CsSO.sub.4, Be(OH).sub.2, BeCO.sub.3, BePO.sub.4, BeO, BeCl.sub.2,
BeBr.sub.2, BeI.sub.2, BeC.sub.2O.sub.4, BeSO.sub.4, Mg(OH).sub.2,
MgCO.sub.3, MgPO.sub.4, MgO, MgCl.sub.2, MgBr.sub.2, MgI.sub.2,
MgC.sub.2O.sub.4, MgSO.sub.4, Ca(OH).sub.2, CaO, CaCO.sub.3,
CaPO.sub.4, CaCl.sub.2, CaBr.sub.2, CaI.sub.2, Ca(OH).sub.2,
CaC.sub.2O.sub.4, CaSO.sub.4, Y.sub.2O.sub.3,
Y.sub.2(CO.sub.3).sub.3, Y.sub.2(PO.sub.4).sub.3, Y(OH).sub.3,
YCl.sub.3, YBr.sub.3, YI.sub.3, Y.sub.2(C.sub.2O4).sub.3,
Y.sub.2(SO4).sub.3, Zr(OH).sub.4, Zr(CO.sub.3).sub.2,
Zr(PO.sub.4).sub.2, ZrO(OH).sub.2, ZrO2, ZrCl.sub.4, ZrBr.sub.4,
ZrI.sub.4, Zr(C.sub.2O.sub.4).sub.2, Zr(SO.sub.4).sub.2,
Ti(OH).sub.4, TiO(OH).sub.2, Ti(CO.sub.3).sub.2,
Ti(PO.sub.4).sub.2, TiO.sub.2, TiCl.sub.4, TiBr.sub.4, TiI.sub.4,
Ti(C.sub.2O.sub.4).sub.2, Ti(SO.sub.4).sub.2, BaO, Ba(OH).sub.2,
BaCO.sub.3, BaPO.sub.4, BaCl.sub.2, BaBr.sub.2, BaI.sub.2,
BaC.sub.2O.sub.4, BaSO.sub.4, La(OH).sub.3,
La.sub.2(CO.sub.3).sub.3, La.sub.2(PO.sub.4).sub.3,
La.sub.2O.sub.3, LaCl.sub.3, LaBr.sub.3, LaI.sub.3,
La.sub.2(C.sub.2O.sub.4).sub.3, La.sub.2(SO.sub.4).sub.3,
Ce(OH).sub.4, Ce(CO.sub.3).sub.2, Ce(PO.sub.4).sub.2, CeO.sub.2,
Ce.sub.2O.sub.3, CeCl.sub.4, CeBr.sub.4, CeI.sub.4,
Ce(C.sub.2O.sub.4).sub.2, Ce(SO.sub.4).sub.2, ThO.sub.2,
Th(CO.sub.3).sub.2, Th(PO.sub.4).sub.2, ThCl.sub.4, ThBr.sub.4,
ThI.sub.4, Th(OH).sub.4, Th(C.sub.2O.sub.4).sub.2,
Th(SO.sub.4).sub.2, Sr(OH).sub.2, SrCO.sub.3, SrPO.sub.4, SrO,
SrCl.sub.2, SrBr.sub.2, SrI.sub.2, SrC.sub.2O.sub.4, SrSO.sub.4,
Sm.sub.2O.sub.3, Sm.sub.2(CO.sub.3).sub.3,
Sm.sub.2(PO.sub.4).sub.3, SmCl.sub.3, SmBr.sub.3, SmI.sub.3,
Sm(OH).sub.3, Sm.sub.2(CO3).sub.3, Sm.sub.2(C.sub.2O.sub.3).sub.3,
Sm.sub.2(SO.sub.4).sub.3, LiCa.sub.2Bi.sub.3O.sub.4Cl.sub.6,
Na.sub.2WO.sub.4, K/SrCoO.sub.3, K/Na/SrCoO.sub.3, Li/SrCoO.sub.3,
SrCoO.sub.3, molybdenum oxides, molybdenum hydroxides, molybdenum
carbonates, molybdenum phosphates, molybdenum chlorides, molybdenum
bromides, molybdenum iodides, molybdenum oxalates, molybdenum
sulfates, manganese oxides, manganese chlorides, manganese
bromides, manganese iodides, manganese hydroxides, manganese
oxalates, manganese sulfates, manganese tungstates, vanadium
oxides, vanadium carbonates, vanadium phosphates, vanadium
chlorides, vanadium bromides, vanadium iodides, vanadium
hydroxides, vanadium oxalates, vanadium sulfates, tungsten oxides,
tungsten carbonates, tungsten phosphates, tungsten chlorides,
tungsten bromides, tungsten iodides, tungsten hydroxides, tungsten
oxalates, tungsten sulfates, neodymium oxides, neodymium
carbonates, neodymium phosphates, neodymium chlorides, neodymium
bromides, neodymium iodides, neodymium hydroxides, neodymium
oxalates, neodymium sulfates, europium oxides, europium carbonates,
europium phosphates, europium chlorides, europium bromides,
europium iodides, europium hydroxides, europium oxalates, europium
sulfates rhenium oxides, rhenium carbonates, rhenium phosphates,
rhenium chlorides, rhenium bromides, rhenium iodides, rhenium
hydroxides, rhenium oxalates, rhenium sulfates, chromium oxides,
chromium carbonates, chromium phosphates, chromium chlorides,
chromium bromides, chromium iodides, chromium hydroxides, chromium
oxalates, chromium sulfates, potassium molybdenum oxides and the
like.
The term "salt," as used herein, generally refers to a compound
comprising negative and positive ions. Salts are generally
comprised of cations and counter ions. Under appropriate
conditions, e.g., the solution also comprises a template, the metal
ion (M.sup.n+) and the anion (X.sup.m-) bind to the template to
induce nucleation and growth of a nanowire of M.sub.mX.sub.n on the
template. "Anion precursor" thus is a compound that comprises an
anion and a cationic counter ion, which allows the anion (X.sup.m-)
to dissociate from the cationic counter ion in a solution. Specific
examples of the metal salt and anion precursors are described in
further detail herein.
The term "oxide," as used herein, generally refers to a metal or
semiconductor compound comprising oxygen. Examples of oxides
include, but are not limited to, metal oxides (M.sub.xO.sub.y),
metal oxyhalides (M.sub.xO.sub.yX.sub.z), metal hydroxyhalides
(M.sub.xOH.sub.yX.sub.z), metal oxynitrates
(M.sub.xO.sub.y(NO.sub.3).sub.z), metal phosphates
(M.sub.x(PO.sub.4).sub.y), metal oxycarbonates
(M.sub.xO.sub.y(CO.sub.3).sub.z), metal carbonates
(M.sub.x(CO.sub.3).sub.z), metal sulfates
(M.sub.x(SO.sub.4).sub.z), metal oxysulfates
(M.sub.xO.sub.y(SO.sub.4).sub.z), metal phosphates
(M.sub.x(PO.sub.4).sub.z), metal acetates
(M.sub.x(CH.sub.3CO.sub.2).sub.z), metal oxalates
(M.sub.x(C.sub.2O.sub.4).sub.z), metal oxyhydroxides
(M.sub.xO.sub.y(OH).sub.z), metal hydroxides (M.sub.x(OH).sub.z),
hydrated metal oxides (M.sub.xO.sub.y).(H.sub.2O).sub.z and the
like, wherein X is independently, at each occurrence, fluoro,
chloro, bromo or iodo, and x, y and z are independently numbers
from 1 to 100.
The term "mixed oxide" or "mixed metal oxide," as used herein,
generally refers to a compound comprising two or more metals and
oxygen (i.e., M1.sub.xM2.sub.yO.sub.z, wherein M1 and M2 are the
same or different metal elements, O is oxygen and x, y and z are
numbers from 1 to 100). A mixed oxide may comprise metal elements
in various oxidation states and may comprise more than one type of
metal element. For example, a mixed oxide of manganese and
magnesium comprises oxidized forms of magnesium and manganese. Each
individual manganese and magnesium atom may or may not have the
same oxidation state. Mixed oxides comprising 2, 3, 4, 5, 6 or more
metal elements can be represented in an analogous manner. Mixed
oxides also include oxy-hydroxides (e.g., M.sub.xO.sub.yOH.sub.z,
wherein M is a metal element, O is oxygen, x, y and z are numbers
from 1 to 100 and OH is hydroxy). Mixed oxides may be represented
herein as M1-M2, wherein M1 and M2 are each independently a metal
element.
The term "crystal domain," as used herein, generally refers to a
continuous region over which a substance is crystalline.
The term "single-crystalline" or "mono-crystalline," as used
herein, generally refers to a material (e.g., nanowire) having a
single crystal domain.
The term "dopant" or "doping agent," as used herein, generally
refers to a material (e.g., impurity) added to or incorporated
within a catalyst to alter (e.g., optimize) catalytic performance
(e.g. increase or decrease catalytic activity). As compared to the
undoped catalyst, a doped catalyst may increase or decrease the
selectivity, conversion, and/or yield of a reaction catalyzed by
the catalyst.
The term "OCM catalyst," as used herein, generally refers to a
catalyst capable of catalyzing an OCM reaction.
"Group 1" elements include lithium (Li), sodium (Na), potassium
(K), rubidium (Rb), cesium (Cs), and francium (Fr).
"Group 2" elements include beryllium (Be), magnesium (Mg), calcium
(Ca), strontium (Sr), barium (Ba), and radium (Ra).
"Group 3" elements include scandium (Sc) and yttrium (Y).
"Group 4" elements include titanium (Ti), zirconium (Zr), halfnium
(Hf), and rutherfordium (Rf).
"Group 5" elements include vanadium (V), niobium (Nb), tantalum
(Ta), and dubnium (Db).
"Group 6" elements include chromium (Cr), molybdenum (Mo), tungsten
(W), and seaborgium (Sg).
"Group 7" elements include manganese (Mn), technetium (Tc), rhenium
(Re), and bohrium (Bh).
"Group 8" elements include iron (Fe), ruthenium (Ru), osmium (Os),
and hassium (Hs).
"Group 9" elements include cobalt (Co), rhodium (Rh), iridium (Ir),
and meitnerium (Mt).
"Group 10" elements include nickel (Ni), palladium (Pd), platinum
(Pt) and darmistadium (Ds).
"Group 11" elements include copper (Cu), silver (Ag), gold (Au),
and roentgenium (Rg).
"Group 12" elements include zinc (Zn), cadmium (Cd), mercury (Hg),
and copernicium (Cn).
"Lanthanides" include lanthanum (La), cerium (Ce), praseodymium
(Pr), neodymium (Nd), promethium (Pm), samarium (Sm), europium
(Eu), gadolinium (Gd), terbium (Tb), dysprosium (Dy), holmium (Ho),
erbium (Er), thulium (Tm), yitterbium (Yb), and lutetium (Lu).
"Actinides" include actinium (Ac), thorium (Th), protactinium (Pa),
uranium (U), neptunium (Np), plutonium (Pu), americium (Am), curium
(Cm), berklelium (Bk), californium (Cf), einsteinium (Es), fermium
(Fm), mendelevium (Md), nobelium (No), and lawrencium (Lr).
"Rare earth" elements include Group 3, lanthanides and
actinides.
"Metal element" or "metal" is any element, except hydrogen,
selected from Groups 1 through 12, lanthanides, actinides, aluminum
(Al), gallium (Ga), indium (In), tin (Sn), thallium (Tl), lead
(Pb), and bismuth (Bi). Metal elements include metal elements in
their elemental form as well as metal elements in an oxidized or
reduced state, for example, when a metal element is combined with
other elements in the form of compounds comprising metal elements.
For example, metal elements can be in the form of hydrates, salts,
oxides, as well as various polymorphs thereof, and the like.
The term "semi-metal element," as used herein, generally refers to
an element selected from boron (B), silicon (Si), germanium (Ge),
arsenic (As), antimony (Sb), tellurium (Te), and polonium (Po).
The term "non-metal element," as used herein, generally refers to
an element selected from carbon (C), nitrogen (N), oxygen (O),
fluorine (F), phosphorus (P), sulfur (S), chlorine (Cl), selenium
(Se), bromine (Br), iodine (I), and astatine (At).
The term "higher hydrocarbon," as used herein, generally refers to
a higher molecular weight and/or higher chain hydrocarbon. A higher
hydrocarbon can have a higher molecular weight and/or carbon
content that is higher or larger relative to starting material in a
given process (e.g., OCM or ETL). A higher hydrocarbon can be a
higher molecular weight and/or chain hydrocarbon product that is
generated in an OCM or ETL process. For example, ethylene is a
higher hydrocarbon product relative to methane in an OCM process.
As another example, a C.sub.3+ hydrocarbon is a higher hydrocarbon
relative to ethylene in an ETL process. As another example, a
C.sub.5+ hydrocarbon is a higher hydrocarbon relative to ethylene
in an ETL process. In some cases, a higher hydrocarbon is a higher
molecular weight hydrocarbon.
The present disclosure is generally directed to processes and
systems for use in the production of hydrocarbon compositions.
These processes and systems may be characterized in that they
derive the hydrocarbon compositions from ethylene that is, in turn,
derived from methane, for example as is present in natural gas. The
disclosed processes and systems are typically further characterized
in that the process for conversion of methane to ethylene is
integrated with one or more processes or systems for converting
ethylene to one or more higher hydrocarbon products, which, in some
embodiments, comprise liquid hydrocarbon compositions. By
converting the methane present in natural gas to a liquid material,
one can eliminate one of the key hurdles involved in exploitation
of the world's vast natural gas reserves, namely transportation. In
particular, exploitation of natural gas resources traditionally has
required extensive, and costly pipeline infrastructures for
movement of gas from the wellhead to its ultimate destination. By
converting that gas to a liquid material, more conventional
transportation systems become available, such as truck, rail car,
tanker ship, and the like.
In some embodiments, processes and systems provided herein include
multiple (i.e., two or more) ethylene conversion process paths
integrated into the overall processes or systems, in order to
produce multiple different higher hydrocarbon compositions from the
single original methane source. Further advantages are gained by
providing the integration of these multiple conversion processes or
systems in a switchable or selectable architecture whereby a
portion or all of the ethylene containing product of the methane to
ethylene conversion system is selectively directed to one or more
different process paths, for example two, three, four, five or more
different process paths to yield as many different products. This
overall process flow is schematically illustrated in FIG. 1. As
shown, an oxidative coupling of methane ("OCM") reactor system 100
is schematically illustrated that includes an OCM reactor train 102
coupled to a OCM product gas separation train 104, such as a
cryogenic separation system. The ethylene rich effluent (shown as
arrow 106) from the separation train 104 is shown being routed to
multiple different ethylene conversion reactor systems and
processes 110, e.g., ethylene conversion systems 110a-110e, which
each produce different hydrocarbon products, e.g., products
120a-120e.
As noted, the fluid connection between the OCM reactor system 100
and each of the different ethylene conversion systems 110a-110e,
can be a controllable and selective connection in some embodiments,
e.g., a valve and control system, that can apportion the output of
the OCM reactor system to one, two, three, four, five or more
different ethylene conversion systems. Valve and piping systems for
accomplishing this may take a variety of different forms, including
valves at each piping junction, multiport valves, multi-valve
manifold assemblies, and the like.
Ethylene-to-Liquids (ETL) Systems
Ethylene-to-liquids (ETL) systems and methods of the present
disclosure can be used to form various products, including
hydrocarbon products. Products and product distributions can be
tailored to a given application, such as products for use as fuel
(e.g., jet fuel or automobile fuels such as diesel or
gasoline).
The present disclosure provides reactors for the conversion of
olefins to higher molecular weight hydrocarbons, which can be in
liquid form. Such reactors can be ETL reactors, which can be used
to convert ethylene and/or other olefins to higher molecular weight
hydrocarbons.
An ETL system (or sub-system) can include one or more reactors. An
ETL system can include at least 1, 2, 3, 4, 5, 6, 7, 8, 9, 10, 11,
12, 13, 14, 15, 16, 17, 18, 19, or 20 ETL reactors, which can be in
a parallel, serial, or a combination of parallel and serial
configuration.
An ETL reactor can be in the form of a tube, packed bed, moving bed
or fluidized bed. An ETL reactor can include a single tube or
multiple tubes, such as a tube in a shell. A multi-tubular reactor
can be used for highly exothermic conversions, such as the
conversion of ethylene to other hydrocarbons. Such a design can
allow for an efficient management of thermal fluxes and the control
of reactor and catalyst bed temperatures.
An ETL reactor can be an isothermal or adiabatic reactor. An ETL
reactor can have one or more of the following: 1) multiple cooling
zones and arrangements within the reactor shell in which the
temperature within each cooling zone may be independently set and
controlled; 2) multiple residence times of the reactants as they
traverse the tubular reactor from the inlet of the individual tubes
to the outlet; and 3) multiple pass design in which the reactants
may make several passes within the reactor shell from the inlet of
the reactor to the outlet.
Multi-tubular reactors of the present disclosure can be used to
convert ethylene to liquid hydrocarbons in a variety of ways. In
some cases, the disclosed multi-tubular ETL reactors can result in
smaller reactors and gas compressors compared to adiabatic ETL
designs. The ETL hydrocarbon reaction is exothermic and thus
reaction heat management may be important for reaction control and
improved product selectivity. In adiabatic ETL reactor designs,
there is an upper limit to the ethylene concentration that is
flowed through reactor due to the amount of heat released and
subsequent temperature rise inside the reactor. To control the heat
of reaction, adiabatic reactors can use a large amount of diluent
gas to mitigate the temperature rise in the reactor bed. In some
cases, the heat of reaction can be managed using multiple reactors
with cooling between reactors and limited conversion between
reactors (i.e., about 20%, about 30%, about 40%, about 50%, about
60%, or about 70% conversion in one reactor), cooling of the
product effluent, and converting the remaining feedstock in one or
more subsequent reactors. The use of diluent gas can result in
larger catalyst beds, reactors, and gas compressors. The
multi-tubular reactors described herein can allow for significantly
greater ethylene concentrations while controlling the reactor bed
temperature, since heat can be removed at the reactor wall. As a
consequence, for a targeted rate of production, smaller catalyst
beds, reactors, and gas compressors may be used.
In addition to smaller ETL reactors, the disclosed multi-tubular
ETL reactors can result in smaller downstream liquid-gas product
separation equipment due to less diluent gas needed to cool the
reactor.
Multi-tubular ETL reactors of the present disclosure can result in
more favorable process conditions toward higher carbon number
hydrocarbon liquids compared to an adiabatic ETL design. Relative
to adiabatic reactors where ethylene feed can be diluted to control
reaction temperature, the disclosed multi-tubular designs can allow
for more concentrated ethylene feed into the reactor while
maintaining good reactor temperature control. Higher ethylene
concentration in the reactor can facilitate the formation for
higher hydrocarbon liquids such as jet and/or diesel fuel since
reactant concentration is important process parameter to yield
higher hydrocarbon oligomers. In some cases, olefinic liquids of
specific carbon number range and types can also be recycled into
the reactor bed to further generate higher carbon number liquids
(e.g., jet/diesel).
Multi-tubular reactors can have multiple temperature zones and
offer multiple residence times. This can allow a wide range of
process flexibility to target a particular product slate. As an
example, a reactor can have multiple temperature zones and/or
residence times. One use of this design can be to make jet and/or
diesel fuel from ethylene. Ethylene oligomerization can require a
relatively high reaction temperature. The temperature required to
react ethylene, to start the oligomerization process may not be
compatible with jet or diesel products, due to the rapid cracking
and/or disproportionation of these jet/diesel products at elevated
temperatures. Multiple reactor temperature zones can allow for a
separate and higher temperature zone to start ethylene
oligomerization while having another lower temperature zone to
facilitate further oligomerization into jet/diesel fuel while
discouraging cracking and disproportionation side reactions.
The use of multiple temperature zones may require different
residence times within a reactor bed. In the jet/diesel example,
the residence time for the ethylene reaction can be different than
the residence time for a lower temperature finishing step to form
jet/diesel. To maximize jet/diesel liquid yield, the ethylene
oligomerization reaction bed temperature may need to be higher but
with a lower residence time than the step to make jet/diesel which
can require a lower temperature but higher residence time. In
adiabatic ETL reactors, multi-temperature processes may occur over
multiple reactor beds with a different temperature associated with
each reactor. The multi-temperature zone approach disclosed herein
can obviate the need for multiple reactors, as in the adiabatic ETL
case, since multiple temperature zones can be achieved within a
single reactor and thus lower capital outlay for reactor
deployment.
Catalyst aging can be an important design constraint in ETL
reaction engineering. ETL catalysts can deactivate over time until
the catalyst bed is no longer able to sustain high ethylene
conversion. A slower catalyst deactivation rate may be desired
since more ethylene can be converted per catalyst bed before the
catalyst bed can need to be taken off-line and regenerated.
Typically, the catalyst deactivates due to "coke", deposits of
carbonaceous material, which results in decreasing catalyst
performance upon coke build-up. The rate of "coke" build-up is
attributable to many different parameters. In ETL adiabatic
reactors, the formation of catalyst bed "hot-spots" can play an
important role in causing catalyst coking. "Hot-spots" favor
aromatic compound formation, which are precursors to coke
formation. "Hot-spots" are a result of temperature non-uniformities
within the catalyst bed due to inadequate heat transfer. The
localized "hot-spots" increase the rate of catalyst
coking/deactivation. The disclosed multi-tubular design can
decrease localized "hot-spots" due to better heat transfer
properties of the multi-tubular design relative to the adiabatic
design. It is anticipated that the decrease in catalyst "hot-spots"
can slow catalyst deactivation.
The product slate of the ETL slate is a result of many factors. An
important factor is the catalyst bed temperature. For example,
higher temperatures catalyst bed temperatures can skew the product
slate, for some catalysts, to aromatic products. In large adiabatic
reactors, controlling "hot spot" formation is challenging and
inhomogeneities in the catalyst bed temperature profiles lead a
wider distribution of products. The proposed multi-tubular design
can significantly reduce catalyst bed temperature
inhomogeneities/"hot spots" due to better heat transfer
characteristics relative to the adiabatic design. As a result, a
narrower product distribution can be more readily achieved than
with adiabatic reactor design. While the multi-tubular design
provides excellent catalyst bed temperature uniformity, catalyst
bed temperature bed uniformity can be further enhanced by injection
of "trim gas" and/or "trim liquid."
The heat capacity of "trim gas" can be used to fine-tune the
catalyst bed to a target temperature. Trim gas composition can be
inert/high heat capacity gas for example: ethane, propane, butane,
and other high heat capacity hydrocarbons.
In some cases, liquid hydrocarbons can be injected into the ETL
reactors to take advantage of the heat of vaporization to further
regulate and cool the reactor bed in order to achieve the desired
temperature. Also, one can use both of them (gases and liquids) as
"trim" agents in this design for ETL.
ETL catalysts may need to be regenerated from a state of low
ethylene conversion (e.g., 20% or less) to high ethylene
conversion, such as, e.g., greater than 20%, 30%, 40%, 50%, 60%, or
70%. Regeneration can occur by heating the catalyst bed to an
appropriate temperature while introducing a portion of diluted air.
The oxygen in air can be used to remove coke by combustion and thus
renew catalyst activity. Too much oxygen can cause uncontrolled
combustion, a highly exothermic process, and the resultant catalyst
bed temperature rise may cause irreversible catalyst damage. As a
consequence, the amount of air that is permitted during adiabatic
reactor regeneration is limited and monitored.
The catalyst regeneration time for an adiabatic reactor can be
largely dictated by the amount of oxygen that can be permitted in
the reactor. The greater heat transfer properties of the disclosed
multi-tubular reactors can permit greater concentrations of oxygen
during catalyst regeneration to hasten catalyst regeneration while
ensuring that the catalyst bed temperature does not reach the point
of irreversible catalyst deactivation.
The present disclosure also provides reactor systems for carrying
out ethylene conversion processes. A number of ethylene conversion
processes can involve exothermic catalytic reactions where
substantial heat is generated by the process. Likewise, for a
number of these catalytic systems, the regeneration processes for
the catalyst materials likewise involve exothermic reactions. As
such, reactor systems for use in these processes can generally be
configured to effectively manage excess thermal energy produced by
the reactions, in order to control the reactor bed temperatures to
most efficiently control the reaction, prevent deleterious
reactions, and prevent catalyst or reactor damage or
destruction.
Tubular reactor configurations that may present high wall surface
area per unit volume of catalyst bed may be used for reactions
where thermal control is desirable or otherwise required, as they
can permit greater thermal transfer out of the reactor. Reactor
systems that include multiple parallel tubular reactors may be used
in carrying out the ethylene conversion processes described herein.
In particular, arrays of parallel tubular reactors each containing
the appropriate catalyst for one or more ethylene conversion
reaction processes may be arrayed with space between them to allow
for the presence of a cooling medium between them. Such cooling
medium may include any cooling medium appropriate for the given
process. For example, the cooling medium may be air, water or other
aqueous coolant formulations, steam, oil, upstream of reaction feed
or for very high temperature reactor systems, molten salt
coolants.
In some cases, reactor systems are provided that include multiple
tubular reactors segmented into one, two, three, four or more
different discrete cooling zones, where each zone is segregated to
contain its own, separately controlled cooling medium. The
temperature of each different cooling zone may be independently
regulated through its respective cooling medium and an associated
temperature control system, e.g., thermally connected heat
exchangers, etc. Such differential control of temperature in
different reactors can be used to differentially control different
catalytic reactions, or reactions that have catalysts of different
age. Likewise, it allows for the real time control of reaction
progress in each reactor, in order to maintain a more uniform
temperature profile across all reactors, and therefore synchronize
catalyst lifetimes, regeneration cycles and replacement cycles.
Differentially cooled tubular reactor systems are schematically
illustrated in FIG. 2. As shown, an overall reactor system 200
includes multiple discrete tubular reactors 202, 204, 206 and 208
contained within a larger reactor housing 210. Within each tubular
reactor is disposed a catalyst bed for carrying out a given
catalytic reaction. The catalyst bed in each tubular reactor may be
the same or it may be different from the catalyst in the other
tubular reactors, e.g., optimized for catalyzing a different
reaction, or for catalyzing the same reaction under different
conditions. As shown, the multiple tubular reactors 202, 206, 208
and 210 share a common manifold 212 for the delivery of reactants
to the reactors. However, each individual tubular reactor or subset
of the tubular reactors may alternatively include a single reactant
delivery conduit or manifold for delivering reactants to that
tubular reactor or subset of reactors, while a separate delivery
conduit or manifold is provided for delivery of the same or
different reactants to the other tubular reactors or subsets of
tubular reactors.
As an alternative or in addition to, the reactor systems used in
conjunction with the olefin (e.g., ethylene) conversion processes
described herein can provide for variability in residence time for
reactants within the catalytic portion of the reactor. Residence
time within a reactor can be varied through the variation of any of
a number of different applied parameters, e.g., increasing or
decreasing flow rates, pressures, catalyst bed lengths, etc.
However, a single reactor system may be provided with variable
residence times, despite sharing a single reactor inlet, by varying
the volume of different reactor tubes or reactor tube portions
within a single reactor unit ("catalyst bed length"). As a result
of varied volumes among reactor tubes or reactor tube portions into
which reactants are being introduced at a given flow rate,
residence times for those reactants within those varied volume
reactor tubes or reactor tube portions, can be consequently
varied.
Variation of reactor volumes may be accomplished through a number
of approaches. By way of example, varied volume may be provided by
including two or more different reactor tubes into which reactants
are introduced at a given flow rate, where the two or more reactor
tubes each have different volumes, e.g., by providing varied
diameters. As can be appreciated, the residence time of gases being
introduced at the same flow rate into two or more different
reactors having different volumes can be different. In particular,
the residence time can be greater in the higher volume reactors and
shorter in the smaller volume reactors. The higher volume within
two different reactors may be provided by providing each reactor
with different diameters. Likewise, one can vary the length of the
reactors catalyst bed, in order to vary the volume of the catalytic
portion.
Alternatively, or additionally, the volume of an individual reactor
tube can be varied by varying the diameter of the reactor along its
length, effectively altering the volume of different segments of
the reactor. Again, in the wider reactor segments, the residence
time of gas being introduced into the reactor tube can be longer in
the wider reactor segments than in the narrower reactor
segments.
Varied volumes can also be provided by routing different inlet
reactant streams to different numbers of similarly sized reactor
conduits or tubes. In particular, reactants, e.g., gases, may be
introduced into a single reactor tube at a given flow rate to yield
a particular residence time within the reactor. In contrast,
reactants introduced at the same flow rate into two or more
parallel reactor tubes can have a much longer residence time within
those reactors.
The above-described approaches to varying residence time within
reactor catalyst beds are illustrated with reference to FIGS. 3-4.
FIG. 3 schematically illustrates a reactor system 300 in which two
or more tubular reactors 302 and 304 are disposed, each having its
own catalyst bed, 306 and 308, respectively, disposed therein. The
two reactors are connected to the same inlet manifold such that the
flow rate of reactants being introduced into each of reactors 302
and 304 are the same. Because reactor 304 has a larger volume
(shown as a wider diameter), the reactants can be retained within
catalyst bed 308 for a longer period. In particular, as shown,
reactor 304 has a larger diameter, resulting in a slower linear
velocity of reactants through the catalyst bed 308, than the
reactants passing through catalyst bed 306. As noted above, one can
similarly increase residence time within the catalyst bed of
reactor 304 by providing a longer reactor. However, such longer
reactor bed may be required to have similar back pressure as a
shorter reactor to ensure reactants are introduced at the same flow
rate as the shorter reactor
FIG. 4 schematically illustrates an alternative approach for
varying reactor volumes in order to vary residence times of
reactants in the catalyst bed. As shown, an individual reactor
unit, e.g., reactor tube 400, can be configured to provide for
differing residence times within different portions of the reactor
tube by varying the diameter of the reactor between reactor segment
404, 406 and 408. In particular, by providing a larger diameter of
the reactor tube in segment 404 and 406, respectively, one can
increase the residence time of reactants moving through these
segments, as the linear velocity of the reactants through such
segments decreases.
The residence time of reactants within reactor systems can be
controlled by varying the diameter of the ETL reactor along the
path of fluid flow. In some cases, the reactor system can include
multiple different reactor tubes, where each reactor tube includes
a catalyst bed disposed therein. Differing residence times may be
employed in catalyzing different catalytic reactions, or catalyzing
the same reactions under differing conditions. In particular, one
may wish to vary residence time of a given set of reactants over a
single catalyst system, in order to catalyze a reaction more
completely, catalyze a different or further reaction, or the like.
Likewise, different reactors within the system may be provided with
different catalyst systems that may benefit from differing
residence times of the reactants within the catalyst bed to
catalyze the same or different reactions from each other.
Alternatively or additionally, residence times of reactants within
catalyst beds may be configured to optimize thermal control within
the overall reactor system. In particular, residence times may be
longer at a zone in the reactor system in which removal of excess
thermal energy is less critical or more easily managed, e.g.,
because the overall reaction has not yet begun generating excessive
heat. In contrast, in other zones of the reactor, e.g., where
removal of excess thermal energy is more difficult due to rapid
exothermic reactivity, the reactor portion may only maintain the
reactants for a much shorter time, by providing a narrower reactor
diameter. As can be appreciated, thermal management becomes easier
due to the shorter period of time that the reactants are present
and reacting to produce heat. Likewise, the reduced volume of a
tubular reactor within a reactor housing also provides for a
greater volume of cooling media, to more efficiently remove thermal
energy.
Systems and methods of the present disclosure can employ fixed bed
reactors. Fixed bed reactors can be adiabatic reactors. Fixed bed
adiabatic ETL reactors can provide for simplicity of the reactor
design. No active external cooling mechanism of the reactor may be
necessary. To control the reactor temperature, profile dilution of
the reactive olefin or other feedstocks (e.g., ethylene, propylene,
butenes, pentenes, etc.) may be necessary. The diluent gas can be
any material that is non-reactive or non-poisonous to the ETL
catalyst but preferably has a high heat capacity to moderate the
temperature rise within the catalyst bed. Examples of diluent gases
include nitrogen (N.sub.2), argon, methane, ethane, propane and
helium. The reactive part of the feedstock can be diluted directly
or diluted indirectly in the reactor by recycling process gas to
dilute the feedstock to an acceptable concentration. Temperature
profile can also be controlled by internal reactor heat exchangers
that can actively control the heat within the catalyst bed.
Catalyst bed temperature control can also be achieved by limiting
feedstock conversion within the catalyst bed. To achieve full
feedstock conversion in this scenario, fixed bed adiabatic reactors
are placed in series with heat exchangers between reactors to
moderate temperature rise reactor over reactor. Partial conversion
occurs in each reactor with inter-stage cooling to achieve the
desired conversion and selectivity for the ETL process.
Since ETL catalysts can deactivate over time through coke
deposition, the fixed bed reactors can be taken off-line and
regenerated, such as by an oxidative or non-oxidative process, as
described elsewhere herein. Once regenerated to full activity the
ETL reactors can be put back on-line to process more feedstock.
Systems and methods of the present disclosure can employ the use of
ETL continuous catalyst regeneration reactors. Continuous catalyst
regeneration reactors (CCRR) can be attractive for processes where
the catalyst deactivates over time and need to be taken off-line to
be regenerated. By regenerating the catalyst in a continuous
fashion less catalyst, fewer reactors for the process as well as
fewer required operations are to regenerate the catalyst. There are
two classes of deployments for CCRR reactors: (1) moving bed
reactors and (2) fluidized bed reactors. In moving bed CCRR design,
the pelletized catalyst bed moves along the reactor length and is
removed and regenerated in a separate vessel. Once the catalyst is
regenerated the catalyst pellets are put back in the ETL conversion
reactor to process more feedstock. The online/regeneration process
can be continuous and can maintain a constant flow of active
catalyst in the ETL reactor. In fluidized bed ETL reactors, ETL
catalyst particles are "fluidized" by a combination of ETL process
gas velocity and catalyst particle weight. During bed fluidization,
the bed expands, swirls, and agitates during reactor operation. The
advantages of an ETL fluidized bed reactor are excellent mixing of
process gas within the reactor, uniform temperature within the
reactor, and the ability to continuously regenerate the coked ETL
catalyst.
Other reactor designs, such as moving bed (MBR), fluidized bed, and
slurry bed reactors can also be employed. An exemplary fluidized
bed reactor 410 is shown in FIG. 4B. A gas inlet stream 411 enters
at the bottom of the reactor and a gas outlet stream 412 exits from
the top of the reactor. Solid particles (e.g., catalyst) enter 413
at one side and exit 414 at another. Within the fluidized bed, gas
bubbles 415 can encounter solid particles 416. The reactor can
comprise a distributor 417 for distributing the gas flow. FIG. 4C
shows additional schematics of exemplary reactor configurations for
co-current moving bed reactors (420), counter-current moving bed
reactors (430), fluidized bed reactors (440), and slurry bed
reactors (450). The moving bed and fluidized bed reactors have
separate gas inlet (421, 431, 434), gas outlet (422, 432, 442),
catalyst inlet (423, 433, 443), and catalyst outlet (424, 434, 444)
configurations. The slurry bed reactor has a combined gas/catalyst
inlet 451 and gas/catalyst outlet 452.
The ETL catalyst can be regenerated with methane or natural gas.
The regeneration stream can have oxygen (O.sub.2) or other
oxidizing agent. The concentration of oxygen in the regeneration
stream can be below the limiting oxygen concentration (LOC), such
that the mixture is not flammable. In some embodiments, the
concentration of O.sub.2 in the regeneration stream is less than
about 6%, less than about 5%, less than about 4%, less than about
3%, less than about 2%, or less than about 1%. In some cases, the
concentration of O.sub.2 in the regeneration stream is between 0%
and about 3%. An advantage of regenerating the ETL catalyst with
methane or natural gas is that, following flowing over the ETL
catalyst for regeneration, the stream can be used in the OCM and/or
ETL process (e.g., the stream can be combusted to provide energy).
The use of methane and/or natural gas to regenerate the ETL
catalyst may not introduce any new components into the process to
achieve catalyst regeneration, which can lead to an efficient use
of materials. In some cases, the use of methane and/or natural gas
makes the economics of the process insensitive, or less dependent
on, the period of time that the ETL catalyst can operate between
regeneration cycles.
Catalysts for the Conversion of Olefins to Liquids
The present invention also provides catalysts and catalyst
compositions for ethylene conversion processes, in accordance with
the processes described herein. In some embodiments, the disclosure
provides modified zeolite catalysts and catalyst compositions for
carrying out a number of desired ethylene conversion reaction
processes. In some cases, provided are impregnated or ion exchanged
zeolite catalysts useful in conversion of ethylene to higher
hydrocarbons, such as gasoline or gasoline blendstocks, diesel
and/or jet fuels, as well as a variety of different aromatic
compounds. For example, where one is using ethylene conversion
processes to convert OCM product gases to gasoline or gasoline
feedstock products or aromatic mixtures, one may employ modified
ZSM catalysts, such as ZSM-5 catalysts modified with Ga, Zn, Al, or
mixtures thereof. In some cases, Ga, Zn and/or Al modified ZSM-5
catalysts are preferred for use in converting ethylene to gasoline
or gasoline feedstocks. Modified catalyst base materials other than
ZSM-5 may also be employed in conjunction with the invention,
including, e.g., Y, ferrierite, mordenite, and additional catalyst
base materials described herein.
In some cases, ZSM catalysts, such as ZSM-5 are modified with Co,
Fe, Ce, or mixtures of these and are used in ethylene conversion
processes using dilute ethylene streams that include both carbon
monoxide and hydrogen components (See, e.g., Choudhary, et al.,
Microporous and Mesoporous Materials 2001, 253-267, which is
incorporated herein by reference). In particular, these catalysts
can be capable of co-oligomerizing the ethylene and H.sub.2 and CO
components into higher hydrocarbons, and mixtures useful as
gasoline, diesel or jet fuel or blendstocks of these. In such
embodiments, a mixed stream that includes dilute or non-dilute
ethylene concentrations along with CO/H.sub.2 gases can be passed
over the catalyst under conditions that cause the
co-oligomerization of both sets of feed components. Use of ZSM
catalysts for conversion of syngas to higher hydrocarbons can be
described in, for example, Li, et al., Energy and Fuels 2008,
22:1897-1901, which is incorporated herein by reference in its
entirety.
The present disclosure provides various catalysts for use in
converting olefins to liquids. Such catalysts can include an active
material on a solid support. The active material can be configured
to catalyze an ETL process to convert olefins to higher molecular
weight hydrocarbons.
ETL reactors of the present disclosure can include various types of
ETL catalysts. In some cases, such catalysts are zeolite and/or
amorphous catalysts. Examples of zeolite catalysts include ZSM-5,
Zeolite Y, Beta zeolite and Mordenite. Examples of amorphous
catalysts include solid phosphoric acid and amorphous aluminum
silicate. Such catalysts can be doped, such as using metallic
and/or semiconductor dopants. Examples of dopants include, without
limitation, Ni, Pd, Pt, Zn, B, Al, Ga, In, Be, Mg, Ca and Sr. Such
dopants can be situated at the surfaces, in the pore structure of
the catalyst and/or bulk regions of such catalysts.
Catalyst can be doped with materials that are selected to effect a
given or predetermined product distribution. For example, a
catalyst doped with Mg or Ca can provide selectivity towards
olefins for use in gasoline. As another example, a catalyst doped
with Zn or Ga (e.g., Zn-doped ZSM-5 or Ga-doped ZSM-5) can provide
selectivity towards aromatics. As another example, a catalyst doped
with Ni (e.g., Ni-doped zeolite Y) can provide selectivity towards
diesel or jet fuel.
Catalysts can be situated on solid supports. Solid supports can be
formed of insulating materials, such as TiOx or AlOx, wherein `x`
is a number greater than zero, or ceramic materials.
Catalyst of the present disclosure can have various cycle lifetimes
(e.g., the average period of time between catalyst regeneration
cycles). In some cases, ETL catalysts can have lifetimes of at
least about 50 hours, 100 hours, 110 hours, 120 hours, 130 hours,
140 hours, 150 hours, 160 hours, 170 hours, 180 hours, 190 hours,
200 hours, 210 hours, 220 hours, 230 hours, 240 hours, 250 hours,
300 hours, 350 hours, or 400 hours. At such cycle lifetimes, olefin
conversion efficiencies less than about 90%, 85%, 80%, 75%, 70%,
65%, or 60% may be observed.
Catalysts of the present disclosure can be regenerated through
various regeneration procedures, as described elsewhere herein.
Such procedures can increase the total lifetimes of catalysts
(e.g., length of time before the catalyst is disposed of). An
example of a catalyst regeneration process is provided in Lubo
Zhou, "BP-UOP Cyclar Process," Handbook of Petroleum Refining
Processes, The McGraw-Hill Companies (2004), pages 2.29-2.38, which
is entirely incorporated herein by reference.
In some embodiments, ETL catalysts can be comprised of base
materials (first active components) and dopants (second active
components). The dopants can be introduced to the base materials
through appropriate methods and procedures, such as vapor or liquid
phase deposition. Dopants can be selected from a variety of
elements, including metallic, non-metallic or amphoteric in forms
of elementary substance, ions or compounds. A few representative
doping elements are Ga, Zn, Al, In, Ni, Mg, B and Ag. Such dopants
can be provided by dopant sources. For example, silver can be
provided by way of AgCl or sputtering. The selection of doping
materials can depend on the target product nature, such as product
distribution. For example, Ga is favorable for aromatics-rich
liquid production while Mg is favorable for aromatics-poor liquid
production.
Base materials can be selected from crystalline zeolite materials,
such as ZSM-5, ZSM-11, ZSM-22, Y, beta, mordenite, L, ferrierite,
MCM-41, SAPO-34, SAPO-11, TS-1, SBA 15 or amorphous porous
materials, such as amorphous silicoaluminate (ASA) and solid
phosphoric acid catalysts. The cations of these materials can be
NH.sub.4.sup.+, H.sup.+ or others. The surface areas of these
materials can be in a range of 1 m.sup.2/g to 10000 m.sup.2/g, 10
m.sup.2/g to 5000 m.sup.2/g, or 100 m.sup.2/g to 1000 m.sup.2/g.
The base materials can be directly used for synthesis or undergo
some chemical treatment, such as desilication (de-Si) or
dealumination (de-Al) to further modify the functionalities of
these materials.
The base materials can be directly used for synthesis or undergo
chemical treatment, such as desilication (de-Si) or dealumination
(de-Al), to get derivatives of the base materials. Such treatment
can improve the catalyst lifetime performance by creating larger
pore volumes, such as pores having diameters greater than or equal
to about 1 nanometer (nm), 2 nm, 3 nm, 4, nm, 5 nm, 10 nm, 20 nm,
30 nm, 40 nm, 50 nm, or 100 nm. In some cases, mesopores having
diameters between about 1 nm and 100 nm, or 2 nm and 50 nm are
created. In some examples, silica or alumina, or a combination of
silica and alumina, can be etched from the base material to make a
larger pore structure in the base catalyst that can enhance
diffusion of reactants and products into the catalyst material.
Pore diameter(s) and volume, in addition to porosity, can be as
determined by adsorption or desorption isotherms (e.g.,
Brunauer-Emmett-Teller (BET) isotherm), such as using the method of
Barrett-Joyner-Halenda (BJH). See Barrett E. P. et al., "The
determination of pore volume and area distributions in porous
substances. I. Computations from nitrogen isotherms," J. Am. Chem.
Soc. 1951. V. 73. P. 373-380. Such method can be used to calculate
material porosity and mesopore volumes, in some cases volumes that
are 3-7 times larger than their original materials. In general, any
changes in catalyst structure, composition and morphology can be
measured by technologies of BET, SEM and TEM, etc.
There are various approaches for doping catalysts. In an example,
the doping components can be added to the base materials and their
derivatives through impregnation, in some cases using incipient
wetness impregnation (IWI), ion exchange or framework substitution
in a zeolite synthesis operation. In some cases, IWI can include i)
mixing a salt solution of the doping component with base material,
for which the amount of salt is calculated based on doping level,
ii) drying the mixture in an oven, and iii) calcining the product
at a certain temperature for a certain time, typically
550-650.degree. C., 6-10 hrs. Ion exchange catalyst synthesis can
include i) mixing a salt solution, which can contain at least 1.5,
2, 3, 4, 5, 6, 7, 8, 9, or 10 times excess amount of the doping
component, with base material, ii) heating the mixture, such as,
for example, at a temperature from about 50.degree. C. to
100.degree. C., 60.degree. C. to 90.degree. C., or 70.degree. C. to
80.degree. C. for a time period of at least about 10 minutes, 30
minutes, 1 hour, 2 hours, 3 hours, 4 hours, 5 hours, 6 hours, 7
hours, 8 hours, 9 hours, 10 hours, 11 hours, or 12 hours, to
conduct a first ion exchange, iii) separating the first ion
exchange mother solution, iv) adding a new salt solution and
repeating ii) and iii) to conduct a second ion exchange, v) washing
the wet solid with deionized water to remove or lower the
concentration of soluble components, vi) drying the raw product,
such as air drying or in an oven, and vii) calcining the raw
product at a temperature from about 450.degree. C. to 800.degree.
C., 500.degree. C. to 750.degree. C., or 550.degree. C. to
650.degree. C. for a time period from about 1 hour to 24 hours, 4
hours to 12 hours, or 6 hours to 10 hours.
In some situations, powder catalysts prepared according to methods
of the present disclosure may need to be formed prior to prepared
in predetermined forms (or form factors) prior to use. In some
examples, the forms can be selected from cylinder extrudates,
rings, trilobe, and pellets. The sizes of the forms can be
determined by reactor size. For example, for a 1''-2'' internal
diameter (ID) reactor, 1.7 mm to 3.0 mm extrudates or equivalent
size for other forms can be used. Larger forms can be used for
different commercial scales (such as 5 mm forms). The ETL reactor
inner diameter (ID) can be any diameter, including ranging from 2
inches to 10 feet, from 1 foot to 6 feet, and from 3 feet to 4
feet. In commercial reactors, the diameters of the catalyst (e.g.,
extrudate) can be greater than about 3 mm, greater than about 4 mm,
greater than about 5 mm, greater than about 7 mm, greater than
about 10 mm, greater than about 15 mm, or greater than about 20 mm.
Binding materials (binder) can be used for forming the catalysts
and improving catalyst particle strength. Various solid materials
that are inert towards olefins (e.g., ethylene), such as Boehmite,
alumina, silicate, Bentonite, or kaolin, can be used as
binders.
A wide range of catalyst:binder ratio can be used, such as, from
about 95:5 to 30:70, or 90:10 to 50:50. In some cases, a ratio of
80:20 is used for bench scale and pilot reactor catalyst synthesis.
For formed catalysts, the crush strengths can be in the range of
about 1 N/mm to 60 N/mm, 5 N/mm to 30 N/mm, or 7 N/mm to 15
N/mm.
Catalysts prepared according to methods of the present disclosure
can be tested for the production of various hydrocarbon products,
such as gasoline and/or aromatics production. In some cases, such
catalysts are tested for the production of both gasoline and
aromatics.
In an example, a short-term test condition for gasoline production
is 300.degree. C., atmospheric pressure, WHSV=0.65 hr.sup.-1,
N.sub.2 50% and C.sub.2H.sub.4 50%, two hour runs. In another
example, a short-term test condition for aromatics production is
450.degree. C., atmospheric pressure, WHSV=1.31 hr.sup.-1, N.sub.2
50% and C.sub.2H.sub.4 50%, two hour runs. In addition to
conducting the two hour short-term test to obtain the initial
catalytic activity data, for some selected catalysts, the long-term
test (lifetime test) are also performed to obtain data of catalyst
lifetime, catalyst capacity as well as average product composition
over the lifetime runs.
In an example, the results on an initial catalytic activity test at
gasoline production conditions is C.sub.2H.sub.4 conversion greater
than about 99%, C.sub.5+ C mole selectivity greater than about 65%
(e.g., 65%-70%), and C.sub.5+ C mole yield greater than about 65%
(e.g., 65%-70%). Catalyst lifetime performance in one cycle run at
gasoline conditions can be at least about 189 hours, cut at
conversion down to 80%; catalyst capacity is about 182
g-C.sub.2H.sub.4 converted per g-catalyst with C mole yield of
C.sub.5++C.sub.3= C.sub.4- greater than about 70%. With recycling,
C.sub.3= and C.sub.4= can be accounted as liquid products.
In another example, the results on an initial catalytic activity at
aromatics production conditions is C.sub.2H.sub.4 conversion
greater than about 99%, C.sub.5.sup.+ C mole selectivity greater
than about 75% (e.g., 75-80%), C.sub.5+ C mole yield greater than
about 75% (e.g., 75-80%) and aromatics in C.sub.5+ greater than
about 90%. Catalyst lifetime performance in one cycle run at
aromatics production conditions can be at least about 228 hours,
cut at conversion down to 82%, catalyst capacity 143
g-C.sub.2H.sub.4 converted/g-catalyst with average C.sub.5+ yield
around 72% and aromatics yield around 62%.
An ETL catalysts can have a porosity that is selected to optimize
catalyst performance, including selectivity, lifetime, and product
output. The porosity of an ETL catalyst can be between about 4
Angstroms to about 1 micrometer, from 0.01 nm to 500 nm, from 0.1
nm to 100 nm, or from 1 nm to 10 nm as measured by pore symmetry
(e.g., nitrogen porosimetry). An ETL catalyst can have a base
material with a set of pores that have an average pore size (e.g.,
diameter) from about 4 Angstroms to 100 nm, or 4 Angstroms to 10
nm, or 4 Angstroms to 10 Angstroms.
The catalytic materials may also be employed in any number of
forms. In this regard, the physical form of the catalytic materials
may contribute to their performance in various catalytic reactions.
In particular, the performance of a number of operating parameters
for a catalytic reactor that impact its performance can be
significantly impacted by the form in which the catalyst is
disposed within the reactor. The catalyst may be provided in the
form of discrete particles, e.g., pellets, extrudates or other
formed aggregate particles, or it may be provided in one or more
monolithic forms, e.g., blocks, honeycombs, foils, lattices, etc.
These operating parameters include, for example, thermal transfer,
flow rate and pressure drop through a reactor bed, catalyst
accessibility, catalyst lifetime, aggregate strength, performance,
and manageability.
In some cases, it is also desirable that the catalyst forms used
will have crush strengths that meet the operating parameters of the
reactor systems. In particular, a catalyst particle crush strength
should generally support both the pressure applied to that particle
from the operating conditions, e.g., gas inlet pressure, as well as
the weight of the catalyst bed. In general, it is desirable that a
catalyst particle have a crush strength that is greater than about
1 N/mm.sup.2, and preferably greater than about 10 N/mm.sup.2, for
example greater than 1 N/mm.sup.2, and preferably greater than 10
N/mm.sup.2. As will be appreciated, crush strength may be increased
through the use of catalyst forms that are more compact, e.g.,
having lower surface to volume ratios. However, adopting such forms
may adversely impact performance. Accordingly, forms are chosen
that provide the above described crush strengths within the desired
activity ranges, pressure drops, etc. Crush strength is also
impacted though use of binder and preparation methods (e.g.,
extrusion or pelleting).
For example, in some embodiments the catalytic materials are in the
form of an extrudate or pellet. Extrudates may be prepared by
passing a semi-solid composition comprising the catalytic materials
through an appropriate orifice or using molding or other
appropriate techniques. Pellets may be prepared by pressing a solid
composition comprising the catalytic materials under pressure in
the die of a tablet press. Other catalytic forms include catalysts
supported or impregnated on a support material or structure. In
general, any support material or structure may be used to support
the active catalyst. The support material or structure may be inert
or have catalytic activity in the reaction of interest. For
example, catalysts may be supported or impregnated on a monolith
support. In some particular embodiments, the active catalyst is
actually supported on the walls of the reactor itself, which may
serve to minimize oxygen concentration at the inner wall or to
promote heat exchange by generating heat of reaction at the reactor
wall exclusively (e.g., an annular reactor in this case and higher
space velocities).
The stability of the catalytic materials is defined as the length
of time a catalytic material will maintain its catalytic
performance without a significant decrease in performance (e.g., a
decrease >20%, >15%, >10%, >5%, or greater than 1% in
hydrocarbon or soot combustion activity). In some embodiments, the
catalytic materials have stability under conditions required for
the hydrocarbon combustion reaction of >1 hr, >5 hrs, >10
hrs, >20 hrs, >50 hrs, >80 hrs, >90 hrs, >100 hrs,
>150 hrs, >200 hrs, >250 hrs, >300 hrs, >350 hrs,
>400 hrs, >450 hrs, >500 hrs, >550 hrs, >600 hrs,
>650 hrs, >700 hrs, >750 hrs, >800 hrs, >850 hrs,
>900 hrs, >950 hrs, >1,000 hrs, >2,000 hrs, >3,000
hrs, >4,000 hrs, >5,000 hrs, >6,000 hrs, >7,000 hrs,
>8,000 hrs, >9,000 hrs, >10,000 hrs, >11,000 hrs,
>12,000 hrs, >13,000 hrs, >14,000 hrs, >15,000 hrs,
>16,000 hrs, >17,000 hrs, >18,000 hrs, >19,000 hrs,
>20,000 hrs, >1 yrs, >2 yrs, >3 yrs, >4 yrs or >5
yrs.
Catalyst Poisoning
Catalysts of the present disclosure can be poisoned during the
course of catalytically generating a given product. ETL catalysts,
for instance, can be poisoned upon generating higher molecular
weight hydrocarbons from olefins (e.g., ethylene). The present
disclosure provides various approaches for avoiding such
poisons.
Alkynes can be oligomerized over ETL catalysts, such as zeolites or
acid catalysts. During alkyne oligomerization, the alkynes can be
rapidly transformed into polyaromatic molecules, precursors to
coke, which can deactivate the catalyst. The selectivity for
acetylene to make coke can deactivate the ETL catalyst at a faster
rate than an alkene and the catalyst may need to be taken off line
to be regenerated. Any molecule containing an alkyne functional
group can deactivate the ETL catalyst at a faster rate than an
alkene group. One example is acetylene, an alkyne produced in small
quantities within the OCM process.
An approach for eliminating alkynes from feedstock to an ETL
catalyst is to convert the alkynes to other material that may not
poison the ETL catalyst. For example, alkynes can be selectively
hydrogenated to make olefins using a variety of transition metal
catalysts without hydrogenating the olefins into alkanes. Examples
of these catalysts are Pd, Fe, Co, Ni, Zn, and Cu containing
catalysts. Such catalysts can be incorporated in or more reactors
upstream of ETL catalysts.
Dienes can be oligomerized over ETL catalysts, such as zeolites or
acid catalysts. However during diene oligomerization, dienes can be
rapidly transformed into polydienes molecules, precursors to coke,
which can deactivate the ETL catalyst. The selectivity for dienes
to make coke can rapidly deactivate the ETL catalyst and the
catalyst may need to be taken off line to be regenerated. Any
molecule containing a diene functional group can rapidly deactivate
the ETL catalyst. An example is butadiene, a diene produced in
small quantities within the OCM process.
An approach for eliminating dienes from feedstock to an ETL
catalyst is to convert the dienes to other material that may not
poison the ETL catalyst. For example, dienes can be selectively
hydrogenated to make olefins using a variety of transition metal
catalysts without hydrogenating the olefins into alkanes. Examples
of these catalysts are Pd, Fe, Co, Ni, Zn, and Cu containing
catalysts.
Bases can react to neutralize the acid functionality that catalyzes
ETL reactions. If enough base reacts with the ETL catalyst, the
catalyst may no longer be active toward oligomerization and may
need to be regenerated. Bases include nitrogen containing
compounds, particularly ammonia, amines, pyridines, pyroles, and
other organic nitrogen containing compounds. Metal hydroxide
compounds such as lithium, sodium, potassium, cesium hydroxides and
group IIA metal hydroxides may deactivate the catalyst as well as
carbonates of group IA and IIA metals.
Bases can be removed from feedstock to an ETL reactor by, for
example, contacting the feedstock stream with water. This can
remove or decrease the concentration of bases, such as amines,
carbonates, and hydroxides.
Sulfur-containing compounds can deactivate ETL catalysts,
particularly if the catalysts are doped with transition metal
compounds. Sulfur can irreversible bind to the catalyst or metal
dopant to deactivate the catalyst toward oligomerization. Organic
sulfur compounds such as thiols, disulfides, thiolethers,
thiophenes and others mercaptan compounds can be detrimental to the
ETL catalyst.
Sulfur-containing compounds can be removed from feedstock to an ETL
reactor by gas scrubbing, such as, for example, amine gas
scrubbing. Amines can react with sulfur compounds (e.g., H.sub.2S)
to remove such compounds from gas streams. Other ways of removing
sulfur compounds are by molecular sieves or hydrotreating. Examples
of approaches for removing sulfur-containing compounds from a gas
stream are provided in Nielsen, Richard B., et al. "Treat LPGs with
amines," Hydrocarbon Process 79 (1997): 49-59, which is entirely
incorporated herein by reference.
The impact that certain non-ethylene gases can have on ETL
catalysts is summarized in Table 1.
TABLE-US-00001 TABLE 1 Impact of non-ethylene gases on ETL catalyst
Feedstock General Catalyst Impact N.sub.2 Inert Methane Inert
CO.sub.2 Inert H.sub.2 Coke suppressant H.sub.2O Coke suppressant
but will deactivate catalyst in large quantities ethane Inert
propylene Oligomerizes to gasoline butylene Oligomerizes to
gasoline acetylene Coke accelerator Dienes Coke accelerator CO
Inert
Catalyst Regeneration
During the life cycle of a catalyst (e.g., ETL catalyst),
carbon-containing material (e.g., petroleum coke) can deposit and
accumulate on the catalyst. Over time, such carbon-containing
material can decrease the activity of the catalyst, and can even
render the catalyst incapable of converting a feedstock to a
product. The catalyst may need to be changed or regenerated. There
are various approaches for regenerating an ETL catalyst, such as
oxidative regeneration and non-oxidative regeneration.
In oxidative regeneration, an oxidizing agent (e.g., O.sub.2) can
be directed over the ETL catalyst at elevated temperatures to
remove or decrease the concentration of the carbon-containing
material deposited on or over the catalyst. This can occur by
combusting the carbon-containing material. In some cases, prior to
subjecting the catalyst to the oxidizing agent, the catalyst can be
purged with and inert gas (e.g., He, Ar or N.sub.2) to remove any
volatile or residual hydrocarbon product on the catalyst surface.
The catalyst can be subsequently exposed to the oxidizing agent. In
some cases, the oxidizing agent is O.sub.2 that can be provided by
air.
In an example oxidative regeneration process, the process
conditions and amount of air (or oxygen) can be predetermined to
limit or control the amount of heat and water generated during the
combustion process of removing the coke. The amount of O.sub.2 can
be limited to no more than 50%, 40%, 30%, 20%, 10%, or 5%
concentration. Air can be diluted with N.sub.2 or another gas that
is inert toward combustion to dilute the concentration to less than
or equal to about 50%, 40%, 30%, 20%, 10% or 5%. Process conditions
can be selected to keep the increase in temperature of the ETL
catalyst less than or equal to about 700.degree. C., 650.degree.
C., 600.degree. C., 550.degree. C., or 500.degree. C. during the
regeneration. This can help prevent catalyst damage during the
regeneration process. Oxidative regeneration reactor inlet
temperatures can range from about 100.degree. C. to 800.degree. C.,
150.degree. C. to 700.degree. C., or 200.degree. C. to 600.degree.
C. Inlet gas temperatures can be ramped from low to high
temperatures to safely control the regeneration process. During
oxidative regeneration, process gas pressures can range from about
1 bar (gauge, or "barg") to 100 barg, 1 barg to 80 barg, or 1 barg
to 50 barg.
In non-oxidative regeneration, hydrogen (H.sub.2) and/or
hydrocarbons can be used to regenerate the catalyst bed to improve
catalyst activity of the ETL catalyst. Hydrogen or hydrocarbon
gases can be directed over the catalyst bed at a temperature from
about 100.degree. C. to 800.degree. C., 150.degree. C. to
600.degree. C., or 200.degree. C. to 500.degree. C. This can aid in
removing or decreasing the concentration of carbon-containing
material from the catalyst.
There are other approaches for reducing the concentration of
catalyst poisons. Acetylene can be a poison at low levels. The
acetylene and, in some cases, methyl acetylene, butadiene,
propadiene and benzene, may need to be removed to some permissible
levels. An approach for decreasing the concentration of acetylene
is to direct the acetylene to a hydrogenation reactor that
hydrogenates the acetylene and butadiene to a mixture of ethylene
and ethane as well as butane and/or butene.
The acetylene can be hydrogenated, for example, prior to being
contacted with the ETL catalyst. The acetylene hydrogenation
reaction can be practiced over a palladium-based catalyst, such as
those used to convert acetylene to ethylene in conventional steam
cracking (e.g., the PRICAT.TM. series including models PD 301/1, PD
308/4, PD 308/6, PD 508/1, PD 408/5, PD 408/7 and PD 608/1, which
are commercially available as tablets or spheres supported on
alumina). A palladium-based catalyst can include one or more
metals, including palladium. In some cases, the acetylene
hydrogenation catalyst is a doped or modified version of a
commercially available catalyst.
However, in some cases, applying an acetylene hydrogenation
catalyst to the OCM process that has been developed or optimized
for another process (e.g., steam cracking separations and
purification processes) can result in operational issues and/or
non-optimized performance. For example, in steam cracking, the
acetylene conversion reactor can either be located on the front end
(prior to cryogenic separations) or back end (after cryogenic
separations) of the process. In steam cracking, these differences
in running front end and back end typically have to do with the
ratio of hydrogen to acetylene present, the ethylene to acetylene
ratio, and the non-ethylene olefin (e.g., butadiene) to acetylene
ratio. All of these factors can impact the catalyst selectivity for
forming ethylene from acetylene, the lifetime and regeneration of
the catalyst, green oil formation, specific process conditions for
the reactor, and additional hydrogen required for the reaction.
These factors are also different between steam cracking versus OCM
and/or ETL processes, therefore, provided herein is an acetylene
hydrogenation catalyst that is designed to be used in an OCM
process.
In OCM and/or ETL implementations, the chemical components going
into the acetylene reactor can be different than for steam
cracking. For example, OCM effluent can include carbon monoxide and
hydrogen. Carbon monoxide can be undesirable because it can compete
with the acetylene for the active sites on the hydrogenation
catalyst and lead to lower activity of the catalyst (i.e., by
occupying those active sites). Hydrogen can be desirable because it
is needed for the hydrogenation reaction, however that hydrogen is
present in the OCM effluent in a certain ratio and adjusting that
ratio can be difficult. Therefore, the catalyst described herein
provides the desired outlet concentrations of acetylene, desired
selectivity of acetylene conversion to ethylene, desired conversion
of acetylene, desired lifetime and desired activity in OCM effluent
gas. As used herein, "OCM effluent gas" generally refers to the
effluent taken directly from an OCM reactor, or having first
undergone any number of further unit operations such as changing
the temperature, the pressure, or performing separations on the OCM
reactor effluent. The OCM effluent gas can have CO, H.sub.2 and
butadiene.
In some embodiments, the catalyst decreases the acetylene
concentration below about 100 parts per million (ppm), below about
80 ppm, below about 60 ppm, below about 40 ppm, below about 20 ppm,
below about 10 ppm, below about 5 ppm, below about 3 ppm, below
about 2 ppm, below about 1 ppm, below about 0.5 ppm, below about
0.3 ppm, below about 0.1 ppm, or below about 0.05 ppm.
The concentration of acetylene can be reached in the presence of
carbon monoxide (CO). In some embodiments, the feed stream entering
the acetylene hydrogenation reactor contains at least about 10%, at
least about 9%, at least about 8%, at least about 7%, at least
about 6%, at least about 5%, at least about 4%, at least about 3%,
at least about 2%, or at least about 1% carbon monoxide.
When used in an OCM and/or ETL process, the acetylene hydrogenation
catalyst can have a lifetime of at least about 6 months, at least
about 1 year, at least about 2 years, at least about 3 years, at
least about 4 years, at least about 5 years, at least about 6
years, at least about 7 years, at least about 8 years, at least
about 9 years, or at least about 10 years.
Another option can be to employ the use of a guard bed in front of
the ETL reactor (or reactor train comprising multiple ETL
reactors). The guard bed can enable the ETL reactor to
preferentially coke out the acetylene. Can guard bed can coke
relatively quickly and may need to be placed in a lead-lag
configuration so that one bed can be regenerated while the other
bed is being operated. The guard bed can contain a catalyst, and in
some cases spent ETL catalyst, to perform preferential coking. The
inlet temperature of guard bed can be lower than the inlet
temperature for ETL, and the space velocity can be higher.
In an example, two guard beds are placed upstream of four or five
parallel ETL reactor beds. The two guard beds are designed in a
lead-lag configuration. The inlet temperature of the guard bed is
about 40.degree. C., about 60.degree. C., about 80.degree. C., or
about 100.degree. C. lower than the inlet to the ETL reactors and
the space velocity is at least about 5.times., at least about
10.times., at least about 20.times. or at least about 50.times.
greater than the space velocity of the ETL reactors. The ETL
reactors are on a schedule where each parallel reactor is
regenerated and decoked every three weeks. But the guard bed is
regenerated and decoked every 36 hours.
Catalyst Activators
The lifetime of a catalyst can be increased using activators.
Activators can be used with catalysts of the present disclosure,
such as ETL catalysts. With the aid of activators of the present
disclosure, the lifetime of a catalyst can be increased by at least
about 10 hours, 20 hours, 30 hours, 40 hours, 50 hours, 100 hours,
110 hours, 120 hours, 130 hours, 140 hours, 150 hours, 160 hours,
170 hours, 180 hours, 190 hours, 200 hours, 210 hours, 220 hours,
230 hours, 240 hours, 250 hours, 300 hours, 350 hours, or 400
hours. Activators of the present disclosure can be used to increase
the lifetime of a catalyst by a factor of at least about 1.1, 1.2,
1.3, 1.4, 1.5, 1.6, 1.7, 1.8, 1.9, 2, 3, 4, 5, 6, 7, 8, 9, or 10 in
relation to situations in which activators are not used. Activators
can be molecules included in the process flow that contacts the
catalyst and/or molecules or elements contained in the catalyst
itself (e.g., dopants). For example, Ga-doped ZSM-5 has an
increased lifetime (cycle lifetime and/or replacement lifetime)
relative to non-doped ZSM-5 (e.g., because the doped catalyst has a
lower selectivity for coke formation).
For example, the addition of water can enhance ETL catalyst
lifetime by suppressing coke formation. Coke formation can be
suppressed by water by reacting with coke to form carbon monoxide
and hydrogen. One of the attractive features from the OCM-ETL
process is that water addition can be optimized to have the maximum
benefit for reducing coke formation in the reactor. Water can be
supplied in a concentration of at least about 1%, 2%, 3%, 4%, 5%,
10%, 15%, 20%, 25%, or 30%. In some situations, the concentration
of water in feedstock into an ETL reactor is from 0% to 30%, or 1%
to 25%.
The addition of hydrogen in a feedstock stream into an ETL reactor
can enhance ETL catalyst lifetime. Hydrogen gas (H.sub.2) can be
directed into an ETL reactor and over an ETL catalyst, which can
reduce the concentration of carbon-containing material (e.g., coke)
that may be present on the catalyst and prohibit the deposition of
carbon-containing material by hydrocracking reactions, for example,
by breaking up larger molecules that may be eventually turned into
coke and decrease catalyst activity.
ETL Processes and Operating Conditions
The present disclosure provides methods for operating ETL reactors
to effect a given or predetermined product distribution or
selectivity. The process conditions can be applied across a single
or plurality of ETL reactors in series and/or parallel.
Hydrocarbon streams into or out of an ETL reactor can include
various other non-hydrocarbon material. In some cases, hydrocarbon
streams can include one or more elements leached from an OCM
catalyst (e.g., La, Nd, Sr, W) or ETL catalyst (e.g., Ga
dopant)
Reactor conditions can be selected to provide a given selectivity
and product distribution. In some cases, for catalyst selectivity
towards aromatics, an ETL reactor can be operated at a temperature
greater than or equal to about 300.degree. C., 350.degree. C.,
400.degree. C., 410.degree. C., 420.degree. C., 430.degree. C.,
440.degree. C., 450.degree. C., or 500.degree. C., and a pressure
greater than or equal to about 250 pounds per square inch (PSI)
(absolute), 200 PSI, 250 PSI, 300 PSI, 350 PSI or 400 PSI. For
catalyst selectivity towards jet or diesel fuel, an ETL reactor can
be operated at a temperature greater than or equal to about
100.degree. C., 150.degree. C., 200.degree. C., 210.degree. C.,
220.degree. C., 230.degree. C., 240.degree. C., 250.degree. C., or
300.degree. C., and a pressure greater than or equal to about 350
PSI, 400 PSI, 450 PSI, or 500 PSI. For catalyst selectivity towards
gasoline, an ETL reactor can be operated at a temperature greater
than or equal to about 200.degree. C., 250.degree. C., 300.degree.
C., 310.degree. C., 320.degree. C., 330.degree. C., 340.degree. C.,
350.degree. C., or 400.degree. C., and a pressure greater than or
equal to about 250 PSI, 300 PSI, 350 PSI, or 400 PSI.
In some cases, the operating conditions of an ETL process are
substantially determined by one or more of the following
parameters: process temperature range, weight-hourly space velocity
(mass flow rate of reactant per mass of solid catalyst), partial
pressure of a reactant at the reactor inlet, concentration of a
reactant at the reactor inlet, and recycle ratio and recycle split.
The reactant can be a (light) olefin--e.g., an olefin that has a
carbon number in the range C2-C7, C2-C6, or C2-C5.
Temperatures used in a gasoline process can be from about 150 to
600.degree. C., 220.degree. C. to 520.degree. C., or 270.degree. C.
to 450.degree. C. Lower temperature can result in insufficient
conversion while higher temperatures can result in excessive coking
and cracking of product. In an example, the WHSV can be between
about 0.5 hr.sup.-1 and 3 hr.sup.-1, partial pressures can be
between about 0.5 bar (absolute) and 3 bar, and concentrations at
the reactor inlet can be between about 2% and 30%. Higher
concentrations can yield difficult-to-manage temperature
excursions, while lower concentrations can make it difficult to
achieve sufficiently high partial pressures and separation of the
products. A process can achieve longer catalyst lifetime and higher
average yields when a portion of the effluent is recycled. The
recycle can be determined by a recycle ratio (e.g., volume of
recycle gas/volume of make-up feed) and the post-reactor
vapor-liquid split which determines the composition of the recycle
stream. There may be several degrees of freedom to the recycle
split, but in some cases the composition of the recycle stream may
be important, which is achieved by post-reactor separation (i.e.,
typical carbon number/boiling point range that is recycled vs. the
carbon number/boiling point ranges that are removed by product
and/or secondary process streams.
To achieve longer average chain lengths and to avoid cracking of
elongated chains such as those found in jet fuel and distillates,
ETL can be performed at reactor operating temperatures from about
150.degree. C. to 500.degree. C., 180.degree. C. to 400.degree. C.,
or 200.degree. C. to 350.degree. C. The slower kinetics may suggest
a lower minimum WHSV of about 0.1 hr.sup.-1. Longer chain lengths
may be favored by high partial pressures, so the upper end for
jet/distillates may be higher than for gasoline, in some cases as
high as about 30 bar (absolute), 20 bar, 15 bar, or 10 bar.
More consistent production of aromatics can be achieved at high
temperature ranges, such as a temperature up to about 200.degree.
C., 250.degree. C., 300.degree. C., 350.degree. C., 400.degree. C.,
450.degree. C., or 500.degree. C. In an adiabatic or even in a
pseudo-isothermal reactor, the ethylene/olefin feed can be diluted
by an inert gas (e.g., N.sub.2, Ar, methane, ethane, propane,
butane or He). The inert gas can serve to moderate the temperature
increase in the reactor bed, and maintain and stabilize contact
time. The olefin concentration at the reactor inlet can be less
than about 50%, 40%, 30%, 20%, or 10%. In some cases, the higher
the molar heat capacity of the diluent, the higher the inlet
concentration of olefins can be to achieve the same temperature
rise.
The following is a list of suitable compounds that may be found in
significant quantities in the process. Such compounds are listed in
the order of increasing heat capacity: nitrogen, carbon dioxide,
methane, ethane, propane, n-butane, iso-butane.
In some cases, a continuous process for making mixtures of
hydrocarbons from (light) olefins by oligomerization comprises
feeding olefinic compounds to a reaction zone of an ETL reactor.
The reactor zone can contain a heterogeneous catalyst. One or more
inert gases can be co-fed to the reactor inlet, making up from
about 50% (volume %) to 99%, 60% to 98%, or 70% to 98% of the
feedstock. The mixture can be comprised at least one of the
following compounds: nitrogen, carbon dioxide, methane, ethane,
propane, n-butane, iso-butane. The process (e.g., ETL reactor)
temperature can be between about 150.degree. C. and 600.degree. C.,
180.degree. C. and 550.degree. C., or 200.degree. C. and
500.degree. C. The partial pressure of olefins in the feed can be
between about 0.1 bar (absolute) to 30 bar, 0.1 bar to 15 bar, or
0.2 bar to 10 bar. The total pressure can be between about 1 bar
(absolute) to 100 bar, 5 bar to 50 bar, or 10 bar to 50 bar. The
weight hourly space velocity can be between about 0.05 hr.sup.-1 to
20 hr.sup.-1, 0.1 hr.sup.-1 to 10 hr.sup.-1, or 0.1 hr.sup.-1 to 5
hr.sup.-1.
An effluent or product stream from an ETL reactor can be
characterized by low water content. For example, an ETL product
stream can comprise less than 60 wt %, 56 wt %, 55 wt %, 50 wt %,
45 wt %, 40 wt %, 39 wt %, 35 wt %, 30 wt %, 25 wt %, 20 wt %, 15
wt %, 10 wt %, 5 wt %, 3 wt %, or 1 wt % water.
In some cases, at least a portion of the reactor effluent is
recycled to the reactor inlet. As an alternative, at most a portion
of the reactor effluent is recycled to the reactor inlet. The
volumetric recycle ratio (i.e., flow rate of the recycle gas stream
divided by flow rate of the make-up gas stream (i.e., fresh feed))
can be between about 0.1 and 30, 0.3 and 20, or 0.5 and 10.
A continuous process for making mixtures of hydrocarbons for use as
gasoline can comprise feeding olefinic compounds to a reaction zone
of an ETL reactor. The ETL reactor can include a catalyst that is
selected for gasoline production, as described elsewhere herein.
The process temperature can be between about 200.degree. C. and
600.degree. C., 250.degree. C. and 500.degree. C., or 300.degree.
C. and 450.degree. C. The partial pressure of olefins in the feed
can be between about 0.1 bar (absolute) to 10 bar, 0.3 bar to 5
bar, or 0.5 bar to 3 bar. The total pressure can be between about 1
bar (absolute) to 100 bar, 5 bar to 50 bar, or 10 bar to 50 bar.
The weight hourly space velocity can be between about 0.1 hr.sup.-1
to 20 hr.sup.-1, 0.3 hr.sup.-1 to 10 hr.sup.-1, or 0.5 hr.sup.-1 to
3 hr.sup.-1.
For products in the distillate range (e.g., C.sub.10+ molecules,
which can exclude gasoline in some cases), the catalyst composition
can be selected as described elsewhere herein. The process
temperature can be between about 100.degree. C. and 600.degree. C.,
150.degree. C. and 500.degree. C., or 200.degree. C. and
375.degree. C. The partial pressure of olefins in the feed can be
between about 0.5 bar (absolute) to 30 bar, 1 bar to 20 bar, or 1.5
bar to 10 bar. The total pressure can be between about 1 bar
(absolute) to 100 bar, 5 bar to 50 bar, or 10 bar to 50 bar. The
weight hourly space velocity can be between about 0.05 hr.sup.-1 to
20 hr.sup.-1, 0.1 hr.sup.-1 to 10 hr.sup.-1, or 0.1 hr.sup.-1 to 1
hr.sup.-1.
For products comprising mixtures of hydrocarbons substantially
comprised of aromatics, the catalyst composition can be selected as
described elsewhere herein. The process temperature can be between
about 200.degree. C. and 800.degree. C., 300.degree. C. and
600.degree. C., or 400.degree. C. and 500.degree. C. The partial
pressure of olefins in the feed can be between about 0.1 bar
(absolute) to 10 bar, 0.3 bar to 5 bar, or 0.5 bar to 3 bar. The
total pressure can be between about 1 bar (absolute) to 100 bar, 5
bar to 50 bar, or 10 bar to 50 bar. The weight hourly space
velocity can be between about 0.05 hr.sup.-1 to 20 hr.sup.-1, 0.1
hr.sup.-1 to 10 hr.sup.-1, or 0.2 hr.sup.-1 to 1 hr.sup.-1.
The ETL process can generate a variety of long-chain hydrocarbons,
including normal and isoparaffins, napthenes, aromatics and
olefins, which may not be present in the feed to the ETL reactor.
The catalyst can deactivate due to the deposition of carbonaceous
deposits ("coke") on the surfaces of the catalyst. As the
deactivation progresses, the conversion of the process changes
until a point is reached when the catalyst can be regenerated.
In some cases, in the early stages of a reaction cycle, the product
distribution can contain large fractions of aromatics and
short-chained alkanes. Later stages can feature increased fractions
of olefins. All stages can feature various amounts isoparaffins,
n-paraffins, naphthenes, aromatics, and olefins, including olefins
other than feed olefins. The change in selectivity with time can be
exploited by separating products. For example, the aromatics-rich
effluent characteristic of the early stages of a reaction cycle may
be readily separated from the effluent of a catalyst bed in a later
stage of its cycle. This can result in high selectivities of
individual products. An example of how the product distribution can
change over time is given in FIG. 5, which is for a Ga-ZSM-5
catalyst.
The ETL process can generate various byproducts, such as
carbon-containing byproducts (e.g., coke) and hydrogen. The
selectivity for coke can be on the order of at least about 1%, 2%,
3%, 4%, or 5% over the course of an ETL process. Hydrogen
production can vary with time, and the amount of hydrogen generated
can be correlated with aromatics production.
In some cases, the time-averaged product of the process can yield a
liquid with a composition that meets the specification of
reformulated gasoline blendstock for oxygen blending (RBOB). In
some cases, RBOB has at least about an 93 octane rating using the
(RON+MON)/2 method, has less than about 1.3 vol % benzene as
measured by ASTM D3606, has less than about 50 vol % aromatics as
measured by ASTM D5769, has less than about 25 vol % olefins as
measured by ASTM D1319 and/or D6550, has less than 80 ppm (wt)
sulfur as measured by ASTM D2622, or any combination thereof. Such
liquid can be employed for use as fuel or other combustion
settings. This liquid can be partially characterized by the content
of aromatics. In some cases, this liquid has an aromatics content
from 10% to 80%, 20% to 70%, or 30% to 60%, and an olefins content
from 1% to 60%, 5% to 40%, or 10% to 30%. Gasoline can comprise
about 60% to 95%, 70% to 90%, or 80-90% of such liquid, with the
remainder in some cases being an alcohol, such as ethanol.
In some situations, an ETL process is used to generate a mixture of
hydrocarbons from light olefin compounds (e.g., ethylene). The
mixture can be liquid at room temperature and atmospheric pressure.
The process can be used to form a mixture of hydrocarbons having a
hydrocarbon content that can be tailored for various uses. For
example, mixtures typically characterized as gasoline or distillate
(e.g., kerosene, diesel) blend stock, or aromatic compounds, can
contribute at least 30%, 40%, 50%, 60%, or 70% by weight to the
final fuel product.
The product selectivity of the ETL process can change with time.
With such changes in selectivity, the product can include varying
distributions of hydrocarbons. Separations units can be used to
generate a product distribution which can be suitable for given end
uses, such as gasoline.
Products of ETL processes of the present disclosure can include
other elements or compounds that may be leached from reactors or
catalysts of the system (e.g., OCM and/or ETL reactors). Examples
of OCM catalysts and the elements comprising the catalyst that can
be leached into the product can be found in U.S. patent application
Ser. No. 13/689,611 or U.S. Provisional Patent Application
61/988,063, each of which is incorporated by reference in its
entirety. Such elements can include transition metals and
lanthanides. Examples include, but are not limited to Mg, La, Nd,
Sr, W, Ga, Al, Ni, Co, Ga, Zn, In, B, Ag, Pd, Pt, Be, Ca, and Sr.
The concentration of such elements or compounds can be at least
about 0.01 parts per billion (ppb), 0.05 ppb, 0.1 ppb, 0.2 ppb, 0.3
ppb, 0.4 ppb, 0.5 ppb, 0.6 ppb, 0.7 ppb, 0.8 ppb, 0.9 ppb, 1 ppb, 5
ppb, 10 ppb, 50 ppb, 100 ppb, 500 ppb, 1 part per million (ppm), 5
ppm, 10 ppm, or 50 ppm as measured by inductively coupled plasma
mass spectrometry (ICPMS).
The composition of ETL products from a system can be consistent
over several cycles of catalyst use and regeneration. A reactor
system can be used and regenerated for at least about 10, 20, 30,
40, 50, 60, 70, 80, 90, or 100 cycles. After a number of
regeneration cycles, the composition of the ETL product stream can
differ from the composition of the first cycle ETL product stream
by no more than about 0.1%, 0.2%, 0.3%, 0.4%, 0.5%, 0.6%, 0.7%,
0.8%, 0.9%, 1%, 2%, 3%, 4%, 5%, 6%, 7%, 8%, 9%, 10%, 11%, 12%, 13%,
14%, 15%, 16%, 17%, 18%, 19%, or 20%.
ETL Process Design
The present disclosure provides various approaches for designing an
ETL process. In the oligomerization of C.sub.2H.sub.4, a range of
hydrocarbons can be formed, including C.sub.2H.sub.6 and CH.sub.4,
as well as H.sub.2. The shorter chain hydrocarbons (e.g., C1-C4)
and hydrogen in the product stream can be separated from the
C.sub.5+ liquid fraction. A fraction of the process stream
containing these lighter molecular weight products can be combined,
or recycled, with incoming C.sub.2H.sub.4 feed stream, as shown in
FIG. 6. In this figure, product stream 605 can be separated, for
instance by a condenser/phase separator 606. The gas stream 608
from the condenser/phase separator can be partially recovered 609
and partially recycled 610 back into the reactor 604 for further
contacting with the catalyst. The OCM reactor effluent 601, which
can have been treated and/or compressed, is then routed to the
treatment unit 602 which may comprise of a water removal unit, or
any other purification unit. The treated ETL feed 603 reacts in the
ETL reactor 604 to generate net liquid product 607 from the
condenser and phase separator unit. The condenser and phase
separator unit also sends a recycle 608 back to the ETL reactor
inlet.
Recycling can have various benefits, such as: 1) further reaction
of shorter chain hydrocarbon products to form higher molecular
weight products, 2) increasing catalyst lifetime, and 3) diluting
the C.sub.2H.sub.4 feed stream to control the reactor process
conditions of reactant concentration and adiabatic temperature
rise.
At the same C.sub.2H.sub.4 WHSV, the conversion of a reactor inlet
stream containing recycle can have a higher yield of liquids
production (C.sub.5+), particularly C.sub.5+ condensable at a
temperature of around 0.degree. C., than that of a reactor inlet
stream without recycle products (see FIGS. 7 and 8). The use of
recycle can also increase catalyst lifetime, as measured by
time-on-stream and grams C.sub.2H.sub.4 converted per grams of
catalyst. Recycle ratios and g liquid condensed at about 0.degree.
C. are shown in Table 2.
FIG. 7A shows liquid phase mass balance for C.sub.2H.sub.4
conversion by using single pass reactor. FIG. 7B shows liquid phase
mass balance for C.sub.2H.sub.4 conversion by using a reactor with
5:1 recycle. The reactors are operated at a pressure of about 30
bar, a weight hourly space velocity (WHSV) of about 0.7 h.sup.-1,
and an inlet temperature of about 350.degree. C. The amounts of
various hydrocarbons produced ranges from 0% to 100% for various
ethylene conversions with paraffins 700, isoparaffins 705, olefins
710, napthenes 715, aromatics 720 and C.sub.12+ compounds 725 being
shown. FIG. 8 is a plot showing increasing C.sub.5+ yield (liquid
condensed at about 0.degree. C.) with increasing recycle. The
reaction conditions included a WHSV=0.27 h.sup.-1; reactor inlet
C.sub.2H.sub.4 mol %=2; T.sub.peakbed=315.degree. C.; total
pressure 300 psi (gauge);
TABLE-US-00002 TABLE 2 Reactor conditions characterized by product
stream data shown in FIG. 8, including recycle ratios, process
inlet C.sub.2H.sub.4 mol %, reactor inlet C.sub.2H.sub.4 mol % and
grams of liquid condensed per grams of C.sub.2H.sub.4 fed. Process
inlet Reactor inlet g liquid condensed Recycle ratio C.sub.2H.sub.4
mol % C.sub.2H.sub.4 mol % per g C.sub.2H.sub.4fed 9:1 20 2 0.76
4:1 10 2 0.66 2:1 6 2 0.54 0:1 2 2 0.15 (no recycle)
In some cases, an inlet feed stream that is diluted with recycle
product stream allows for a smaller adiabatic temperature rise in
the reactor and reduced C.sub.2H.sub.4 concentration into the
reactor. A lower adiabatic temperature rise, and therefore peak
reactor temperature, can alter the effluent product stream
composition. Higher peak reactor temperatures, for instance, can
increase the yield and selectivity of aromatic products.
Different amounts of ethylene in an ETL product stream can be
recycled. In some cases, at least about 5%, 10%, 15%, 20%, 25%,
30%, 25%, 40%, 45%, 50%, 55%, 60%, 65%, 70%, 75%, 80%, 85%, 90%,
95%, 96%, 97%, 98%, 99%, or 100% of ethylene in an ETL product
stream is recycled. In some cases, at most about 5%, 10%, 15%, 20%,
25%, 30%, 25%, 40%, 45%, 50%, 55%, 60%, 65%, 70%, 75%, 80%, 85%,
90%, 95%, 96%, 97%, 98%, 99%, or 100% of ethylene in an ETL product
stream is recycled.
An ETL process can be characterized by a single pass conversion or
single pass conversion of C.sub.2+ compounds to C.sub.3+ compounds
of at least 10%, 15%, 20%, 25%, 30%, 35%, 40%, 45%, 50%, 55%, 60%,
65%, 70%, 75%, 80%, 85%, 90%, 95%, 96%, 97%, 98%, 99%, 99.9%, or
99.99%.
ETL Process Feedstock
The feedstock to an ETL reactor can have an effect on the product
distribution out of the ETL reactor. The product distribution can
be related to the concentration of olefins into the ETL reactor,
such as ethylene, propylene, butene(s) and pentene(s). The
feedstock concentration can impact ETL catalyst efficiency. A
feedstock having an olefin concentration that is greater than or
equal to about 5%, 10%, 15%, 20%, 25%, 30%, or 40% can be efficient
at generating higher molecular weight hydrocarbons. In some cases,
the optimum olefin concentration can be less than about 80%, 85%,
75%, 70% or 60%. The ETL feedstock can be characterized based on
the ethylene to ethane molar ratio of the feedstock, which can be
at least about 2:1, 3:1, 4:1, 5:1, 6:1, 7:1, or 8:1.
The presence of other C.sub.2+ compounds and non-C.sub.2+
impurities (e.g., CO, CO.sub.2, H.sub.2O and H.sub.2) can have an
impact on ETL selectivity and/or product distribution. For
instance, the presence of acetylene and/or dienes in a feedstock to
an ETL reactor can have a significant impact on ETL selectivity
and/or product distribution, since acetylene may be a deactivator
and coke accelerator.
Separations for ETL
Separations for ETL processes of the present disclosure can be
carried out in three places within the ETL scheme: before the ETL
reactor, within the ETL reactor and downstream of the ETL reactor.
In each of these three places, different separations technologies
can be employed.
To process the ETL reactor feed, traditional gas separations
equipment can be used. These separations may include pressure swing
adsorption, temperature swing adsorption and membrane-based
separation. The reactor feed could also be augmented by utilizing
cryogenic separations equipment found in a traditional midstream
gas plant.
To make changes to the composition within the reactor, different
types of catalyst can be co-mixed or layered within the catalyst
bed or reactor vessel. Different types of zeolite catalysts (for
example a ZSM-5 and a SAPO 34 in a 60%/40% mixture or in a 50%/50%
mixture) could create different hydrocarbon profiles at the reactor
vessel outlet. Also within this vessel, there could be a
combination of multiple beds with appropriate quenches built in to
affect the final product composition.
To separate the reactor outlet mixtures, a combination of flash
separation, hydrogenation, isomerization and distillation can be
used. Flash separation will remove most of the light fractions of
the hydrocarbon liquid product. This can affect product qualities
like Reid Vapor Pressure. Hydrogenation, isomerization and
distillation can then be used, much like traditional refining
processes, to create a fungible product.
ETL separation can be implemented upstream of an ETL reactor.
Membranes used in conjunction with the ETL process can be used on
the process feedstock to enrich components prior to directing the
feedstock to the ETL reactor. Ethylene may be a component that can
be enriched. Other components of the feedstock may also be
enriched, such as H.sub.2 and/or CO.sub.2. In some cases, CO may be
rejected.
For example, CO in the feedstock may be a catalyst poison. CO can
be removed prior to directing the feedstock to the ETL reactor.
Hydrogen may be an advantageous species to have in the feedstock
because it can reduce coking rates, thus lengthening on-stream time
between de-coke cycles.
In some cases, a membrane separation unit upstream of an ETL
reactor may be employed. The membrane unit can remove at least
about 20%, 30%, 40%, 50% or 60% of one component, or increase the
amount of ethylene from at least about 1%, 2%, 3%, 4% or 5% to at
least about 10%, 15%, 20%, 30%, or 40%.
As another example, ethylene can be enriched using a membrane that
has a certain chemical affinity to ethylene. For oxygen separations
membranes, cobalt can be used within the membranes to chemically
pull oxygen through the membranes. Chemically-modified membranes
can be used to effect such separation.
Another technique that can be employed for upstream separation is
pressure swing adsorption (PSA). Pressure swing adsorption can be
used to remove substantially all of a certain poison, or enrich
ethylene to near purity. In some cases, PSA may be used in place
of, or in addition, membrane. The PSA unit can include at least 2,
3, 4, 5, 6, 7, 8, 9, or 10 vessels that contain an adsorbent. This
adsorbent may be a combination of zeolites, molecular sieves or
activated carbon, for example. Each vessel can contain one or more
adsorbents co-mixed or layered within the vessel. An example of a
PSA unit is shown in FIG. 9. The system shown in FIG. 9 is an
activated carbon based system 900 to separate oxygen from nitrogen.
The system comprises carbon molecular sieves 901 which receive gas
streams 902 to be treated. Part of the gas streams can be released
as off gas 903. Treated gas 904 can comprise produced nitrogen. The
activated carbon system can comprise carbon particles 905, which
can interact with N.sub.2 molecules 906 and O.sub.2 molecules 907.
The carbon particles can have sizes for example from about 1 to
about 20 micrometers, with a pore size from about 0.4 nanometers to
about 25 nanometers.
The PSA units can operate at ETL reactor pressures (e.g., 5-50 bar)
and blow down to atmospheric pressure. Activated carbon, 3A, 4A, 5A
molecular sieves and zeolites can be used in these beds. The
vessels can be operated such that the wanted gases (e.g., ethylene)
pass through the beds at high pressure, and unwanted gases (e.g.,
CO, CO.sub.2 or methane) are blown down out of the bed at low
pressure.
As an example, the specific choice of sorbent can determine the
species that passes through at high pressure or is exhausted at low
pressure. In some cases, a PSA can use layered sorbents, such as to
effect methane and nitrogen separation. Such layering within the
bed allows methane to be the blow down gas, rather than
nitrogen.
PSA technology can also be used in other situations. Multiple beds
can be used in series to further enrich the wanted process gases.
PSA units with at least 2, 3, 4, 5, 6, 7, 8, 9, 10, 20, or 30
vessels may be employed. The PSA can be operated at high
frequencies, which can further promote better separation.
Another separation technique that can be employed for use with ETL
is temperature swing adsorption (TSA). In TSA, temperature changes
are used to effect separation. TSA can be used to separation
hydrocarbons mixtures after the ETL reactor. When gas mixtures are
close to changing phases, TSA can be helpful in removing the heavy
fraction from the light fraction.
The present disclosure also provides in-reactor separations
(product augmentation) approaches. Some of the separations goals
can be achieved within the catalyst bed, or within the reactor
vessel itself, using reactive separations, for example. In reactive
separation, a first molecule can be reacted to form a larger or
smaller molecule that may be separated from a given stream.
In some cases, gas phase ethylene can be condensed to a liquid via
reaction. This augmentation can take two forms within the catalyst
bed: it can augment the product to bring it to within a given
specification, or it can augment the product to remove downstream
equipment. As an example of bringing products into specification, a
hydrogenation catalyst can be co-mixed or layered within the bed,
or as a second bed within a reactor vessel. This catalyst can
utilize the available hydrogen to decrease the olefin content of
the final product. Since fungible gasoline (and many other
products) can have an olefin specification to prevent gumming, this
in situ separation can remove a large amount of olefin content from
the resulting liquid, bringing it to within a given
specification.
A co-mixed bed with multiple types of different zeolite can affect
the overall product composition. For example, a low-aromatic
producing catalyst can be added in an 80%/20% mixture to a typical
ETL catalyst. The resulting product stream can be lower in
aromatics, and can bring an off-spec product to within a given
specification.
As another approach, a downstream (in vessel) isomerization bed can
be used to remove unwanted isomers, like durene. Hydrocarbon
compounds of any appropriate carbon number, such as hydrocarbon
compounds with four or more carbon atoms (C.sub.4+ compounds), can
be isomerized. If a downstream unit is necessary to isomerize
components like durene, or remove components, such as high boiling
point components, an in-bed reactor approach can be employed.
In some situations, a mixture of zeolites that have been augmented
via a process may also provide for a desirable separation. Such
mixture can be used to provide for product augmentation.
The present disclosure also provides separations approaches
downstream of an ETL reactor. Downstream separations equipment for
an ETL process can be similar to equipment employed for use in
refineries. In some cases, downstream unit operations can include
flash separation, isomerization, hydrogenation and distillation,
which can aid in bringing the final product to within a given
specification.
Isomerization equipment can convert unwanted iso-durene into a more
volatile form. Hydrogenation equipment can reduce the amount of
olefins/aromatics in the final product. Distillation can separate
material on the basis of boiling point. These units can be readily
used to create a product having a product distribution as
desired.
Isomerization equipment can be used to upgrade the octane rating of
a hydrocarbon product composition. For example, n-hexane can be
isomerized to i-hexane. N-pentane (62 octane) can be isomerized to
2-methyl-butane (93 octane). Hexane (25 octane) can be isomerized
to 2-methyl-pentane (73 octane).
Alkylation and dimerization units can upgrade lighter fractions,
such as butanes, into more valuable, higher octane products. If the
ETL reactor produces a large amount of butenes compared to butanes,
then dimerization can be used to convert the butene into
isooctene/isooctane.
A catalytic reformer unit can upgrade light naphtha fraction to a
reformate. This unit works by combining molecules and producing
hydrogen. If well-placed, the hydrogen produced in this unit can be
utilized in a downstream unit.
Depending on the size and scale of the ETL reactor, vacuum
distillation can be employed to further refine the hydrocarbon
product outputted by the ETL reactor. If such products are valuable
as lubricants, oils and waxes, then the extra step to vacuum
distill these products can be advantageous. In some cases, the
amount of heavy components produced in the ETL reactor is less than
20%, 15%, 10%, 5% or 1%, but the value generated out of those
products can be substantial.
Another approach for separating hydrocarbons is cryogenic
separation. Such separation can be used to capture C4 and C.sub.5+
compounds from an ETL reactor effluent product stream. In some
cases, a cryogenic separation unit can include a cold box that may
not use traditional deep cryogenic temperatures and may not require
traditional unit operations of demethanizer and deethanizer. Such
cryogenic separation unit may not produce high purity methane,
ethane, or propane products. However, it may produce a mixed (in
some cases primarily methane) stream with impurity ethane, propane,
other light hydrocarbons and inert gases that are acceptable for
use in other settings, such as reinjection to pipeline gas, as
residue gas, or used to meet fuel requirements for power plants or
feedstocks for syngas plants for the production of methanol or
ammonia.
In some examples, a cryogenic separation unit can operate at a
temperature from about -100.degree. C. to -20.degree. C.,
-90.degree. C. to -40.degree. C., or -80.degree. C. to -50.degree.
C. Such temperatures can be obtained through methods that use the
turboexpansion of high pressure pipeline natural gas or
turboexpansion of moderate pressure high methane content feedstock
gas, which may be typical of OCM reactor inlet requirements where
additional cooling may be accomplished using traditional process
plant refrigeration loops, including propane refrigeration or other
mixed refrigerants.
In some cases, there may be substantial recovery of
pressure-reduced power by coupling of turboexpander and residue gas
compressors depending on final destination and usage of lighter
nonreacted and unrecoverable hydrocarbons and other components.
In an example OCM-ETL system, gas is expanded and/or additional
refrigeration cooled and fed to a cryogenic cold box unit, where
heat is exchanged with multiple downstream product streams. It can
then be fed to an OCM reaction and heat recovery section. Pressure
can be increased through multiple process gas compressors, then
heated for ETL and then ETL reaction section. Unrefrigerated
liquids recovery can be accomplished using air and cooling water
utilities before the product gas enters the cryogenic cold box
unit, where it is cooled, pressure reduced for cooling effects, and
additional condensed liquids removed via a liquid-liquid separator.
Separated liquids can reenter the cryogenic cold box unit, where
they are heat exchanged prior to being fed to a depropanizer unit
which removes impurity propane and other light compounds from final
C.sub.4+ product. Separated gas from the liquid-liquid separator
also renter the cryogenic cold box unit where they are heat
exchanged prior to being mixed with depropanizer overhead product
gas and then fed to residue gas compressors based on final residue
gas users. The depropanizer reflux condensation is also provided by
sending this gas stream through the cryogenic cold box unit.
In some cases, a debutanizer column can be installed with bottoms
product from depropanizer as feed. Its use can be to provide RVP
control of final C.sub.4+ product. In some cases, RVP control may
be precluded, other purifications or chemical conversions may be
employed.
ETL Reactor Feedstock
Olefin-to-liquids (e.g., ETL) processes of the present disclosure
can be performed using feedstocks comprising one or more olefins,
such as pure ethylene or diluted ethylene. Ethylene can be mixed
with non-hydrocarbon molecules or other hydrocarbons, including
olefins, paraffins, naphthenes, and aromatics. When a feedstock
comprising these materials is directed over an ETL catalyst, such
as a zeolite catalyst bed at temperatures of at least about
150.degree. C., 200.degree. C., 250.degree. C., or 300.degree. C.,
the reactants can oligomerize to form a combination of longer chain
isomers of olefins and paraffins, naphthenes, and aromatics. The
product slate can include hydrocarbons with carbon numbers between
1 and 19 (i.e., C.sub.1-C.sub.19).
The concentration of ethylene (or other olefin(s)) can be changed
by adjusting the partial pressure of ethylene (or other olefin(s))
at constant total pressure by dilution with an inert gas, such as
nitrogen or methane, or by adding an inert gas to increase the
total pressure while keeping the partial pressure of ethylene
constant. A change in concentration due to changes in the total
pressure may not lead to significant variations in the process
unless the system is operated in an adiabatic mode, in which
temperature spikes introduce additional variability.
In an isothermal reactor operation, a change in concentration via
adjustments in the partial pressure of ethylene can prompt
increases in liquid content and reduction of olefins at the benefit
of paraffins and aromatics. The changes observed in product slate
and liquid formation can depend on the temperature regime and the
class of molecules formed in that regime (i.e., isoparaffins and
aromatics at temperatures below or above about 400.degree. C.,
respectively). For example, increasing the concentration of
ethylene from 5% to 15% at a constant total pressure of 1 bar and a
WHSV of 1 g ethylene/g catalyst/hour can result in a change from
15% to 45% liquids at 300.degree. C.
As the temperature increases, the starting liquid percent
increases, yet the net change upon an increase in concentration
diminishes. For example, at 390.degree. C., increasing the
concentration of ethylene from 5% to 15% at a constant total
pressure of 1 bar can result in a change of 45% to 65% liquids. The
composition of the product can also change with increasing
concentration of ethylene. The trend is uniform with temperature:
as the concentration increases, the content of olefins decreases at
the benefit of paraffin isomers, naphthenes, and aromatics. As the
temperature is increased to at least about 300.degree. C.,
350.degree. C., 400.degree. C. or 450.degree. C. and the product
slate is heavily aromatic, changes in the partial pressure of
ethylene may not change the product slate but can cause a decrease
in the liquid content.
In an adiabatic operation, the concentration of ethylene may result
in a change in the liquid and product slate, which is coupled to
the variations in temperature zones across the reactor bed. In this
mode, the rate of heat transfer from a differential volume unit of
the reactor bed is a function of the heat capacity of the catalyst
and gaseous molecules in the stream--in particular the inert
species. Thus, decreasing the concentration of ethylene helps
increase the heat dissipation and the temperature in the volume
unit. In general, as the concentration of ethylene is increased,
the temperature in the bed can increase and the content of
aromatics and net liquids can also increase at the expense of
paraffins, isoparaffins, olefins, and naphthenes. When the
temperature reaches at least about 300.degree. C., 350.degree. C.,
400.degree. C. or 450.degree. C., the net amount of liquid can
decrease as cracking of the liquid molecules becomes more
prevalent.
In some cases, the addition of other hydrocarbons from a recycle,
refinery or midstream operation combined with the ethylene
feedstock may have a positive effect on the formation of liquids.
The ETL process is an oligomerization reaction, in which
hydrocarbons are combined to form longer chain hydrocarbons. Thus,
introducing hydrocarbons with C.sub.3+ olefin chain length in
addition to the C.sub.2 ethylene promotes the formation of liquid.
As long as the reaction conditions or inherent nature of the
catalyst itself precludes cracking (.beta.-scission) of the
hydrocarbon, the addition of longer chain hydrocarbons in the feed
may yield an oligomerized product that is the sum of the two
molecules. In other words, the barrier to producing longer chain
molecules is reduced by minimizing the number of molecular units at
the start of the reactor (C.sub.2+C.sub.2+C.sub.2+C.sub.2=C.sub.8
vs. C.sub.2+C.sub.6=C.sub.8).
Gas molecules that can be co-fed with ethylene can come from a
recycle stream, natural gas liquids, midstream operations, or
refinery effluents comprising ethane, propylene, propane, butene
isomers, and butane isomers, and other C.sub.4+ olefins. The
general product slate can be more or less unchanged by introducing
propylene, isobutene, and trans-2-butene (with similar expectations
for other butene isomers). At a constant volumetric flowrate of
hydrocarbon species, substitution of a longer chain hydrocarbon for
a shorter chain hydrocarbon (e.g., propylene replacing ethylene)
can result in a higher content of liquid formed.
For example, at T=300.degree. C. with 0.15 bar partial pressure of
hydrocarbon, 1 bar total pressure, a 50:50 mixture of propylene or
isobutene with ethylene increases the liquid yield by 10%-20% in
comparison to a pure ethylene feedstock (an increase in liquids can
be due to an increase in liquid (C.sub.5+) isoparaffins). When the
temperature is 390.degree. C. or higher and aromatic molecules are
the dominant product species, the impact of hydrocarbon length has
less effect on the liquid formation. Regardless, we have found that
the presence of propylene or isobutene in the feed promotes the
formation of liquids (aromatics) to an extent (a few percentage
points) that is greater than using an isolated pure feeds.
Additional paraffins (e.g., ethane, propane, and butane) can
influence may impact an ETL reaction and product distribution. The
introduction of n-paraffins may yield an increase in isoparaffin
content due to isomerization of the molecules on the acid zeolite
catalyst. As the temperature and rate of dehydrogenation increases,
the impact of introduced paraffins may mirror the behavior observed
by adding olefins. Co-feeding C.sub.5+ hydrocarbons with ethylene
may also improve the liquid conversion performance of the ETL
process due to the nature of the oligomerization process.
Oxidative Coupling of Methane (OCM) Processes
In an OCM process, methane (CH.sub.4) reacts with an oxidizing
agent over a catalyst bed to generate C.sub.2+ compounds. For
example, methane can react with oxygen over a suitable catalyst to
generate ethylene, e.g., 2 CH.sub.4+O.sub.2.fwdarw.C.sub.2H.sub.4+2
H.sub.2O (See, e.g., Zhang, Q., Journal of Natural Gas Chem.,
12:81, 2003; Olah, G. "Hydrocarbon Chemistry", Ed. 2, John Wiley
& Sons (2003)). This reaction is exothermic (.DELTA.H=-67
kcals/mole) and has typically been shown to occur at very high
temperatures (>700.degree. C.). Experimental evidence suggests
that free radical chemistry is involved. (Lunsford, J. Chem. Soc.,
Chem. Comm., 1991; H. Lunsford, Angew. Chem., Int. Ed. Engl.,
34:970, 1995). In the reaction, methane (CH.sub.4) is activated on
the catalyst surface, forming methyl radicals which then couples in
the gas phase to form ethane (C.sub.2H.sub.6), followed by
dehydrogenation to ethylene (C.sub.2H.sub.4). Several catalysts
have shown activity for OCM, including various forms of iron oxide,
V.sub.2O.sub.5, MoO.sub.3, Co.sub.3O.sub.4, Pt--Rh, Li/ZrO.sub.2,
Ag--Au, Au/Co.sub.3O.sub.4, Co/Mn, CeO.sub.2, MgO, La.sub.2O.sub.3,
Mn.sub.3O.sub.4, Na.sub.2WO.sub.4, MnO, ZnO, and combinations
thereof, on various supports. A number of doping elements have also
proven to be useful in combination with the above catalysts.
Since the OCM reaction was first reported over thirty years ago, it
has been the target of intense scientific and commercial interest,
but the fundamental limitations of the conventional approach to
C--H bond activation appear to limit the yield of this attractive
reaction under practical operating conditions. Specifically,
numerous publications from industrial and academic labs have
consistently demonstrated characteristic performance of high
selectivity at low conversion of methane, or low selectivity at
high conversion (J. A. Labinger, Cat. Lett., 1:371, 1988). Limited
by this conversion/selectivity threshold, no OCM catalyst has been
able to exceed 20-25% combined C.sub.2 yield (i.e. ethane and
ethylene), and more importantly, all such reported yields operate
at extremely high temperatures (>800.degree. C.). Catalysts and
processes have been described for use in performing OCM in the
production of ethylene from methane at substantially more
practicable temperatures, pressures and catalyst activities. These
are described in U.S. Patent Publication Nos. 2012/0041246,
2013/0023079, and 2013/165728, and U.S. patent application Ser.
Nos. 13/936,783 and 13/936,870 (both filed Jul. 8, 2013), the full
disclosures of each of which is incorporated herein by reference in
its entirety for all purposes.
An OCM reactor can include a catalyst that facilitates an OCM
process. The catalyst may include a compound including at least one
of an alkali metal, an alkaline earth metal, a transition metal,
and a rare-earth metal. The catalyst may be in the form of a
honeycomb, packed bed, or fluidized bed. In some embodiments, at
least a portion of the OCM catalyst in at least a portion of the
OCM reactor can include one or more OCM catalysts and/or
nanostructure-based OCM catalyst compositions, forms and
formulations described in, for example, U.S. Patent Publication
Nos. 2012/0041246, 2013/0023709, 2013/0158322, 2013/0165728, and
pending U.S. application Ser. No. 13/901,309 (filed May 23, 2013)
and Ser. No. 14/212,435 (filed Mar. 14, 2014), each of which is
entirely incorporated herein by reference. Using one or more
nanostructure-based OCM catalysts within the OCM reactor, the
selectivity of the catalyst in converting methane to desirable
C.sub.2+ compounds can be about 10% or greater; about 20% or
greater; about 30% or greater; about 40% or greater; about 50% or
greater; about 60% or greater; about 65% or greater; about 70% or
greater; about 75% or greater; about 80% or greater; or about 90%
or greater.
An OCM reactor can be sized, shaped, configured, and/or selected
based upon the need to dissipate the heat generated by the OCM
reaction. In some embodiments, multiple, tubular, fixed bed
reactors can be arranged in parallel to facilitate heat removal. At
least a portion of the heat generated within the OCM reactor can be
recovered, for example the heat can be used to generate high
temperature and/or pressure steam. Where co-located with processes
requiring a heat input, at least a portion of the heat generated
within the OCM reactor may be transferred, for example, using a
heat transfer fluid, to the co-located processes. Where no
additional use exists for the heat generated within the OCM
reactor, the heat can be released to the environment, for example,
using a cooling tower or similar evaporative cooling device. In
some embodiments, an adiabatic fixed bed reactor system can be used
and the subsequent heat can be utilized directly to convert or
crack alkanes into olefins. In some embodiments, a fluidized bed
reactor system can be utilized. OCM reactor systems useful in the
context of the present invention may include those described in,
for example, U.S. patent application Ser. No. 13/900,898 (filed May
23, 2013), which is incorporated herein by reference in its
entirety for all purposes.
The methane feedstock for an OCM reactor can be provided from
various sources, such as non-OCM processes. In an example, methane
is provided through natural gas, such as methane generated in a
natural gas liquids (NGL) system.
Methane can be combined with a recycle stream from downstream
separation units prior to or during introduction into an OCM
reactor. In the OCM reactor, methane can catalytically react with
an oxidizing agent to yield C.sub.2+ compounds. The oxidizing agent
can be oxygen (O.sub.2), which may be provided by way of air or
enriched air. Oxygen can be extracted from air, for example, in a
cryogenic air separation unit.
To carry out an OCM reaction in conjunction with preferable
catalytic systems, the methane and oxygen containing gases
generally need to be brought up to appropriate reaction
temperatures, e.g., typically in excess of 450.degree. C. for
preferred catalytic OCM processes, before being introduced to the
catalyst, in order to allow initiation of the OCM reaction. Once
that reaction begins or "lights off," then the heat of the reaction
is typically sufficient to maintain the reactor temperature at
appropriate levels. Additionally, these processes may operate at a
pressure above atmospheric pressure, such as in the range of about
1 to 30 bars (absolute).
In some cases, the oxidizing agent and/or methane are
pre-conditioned prior to, or during, the OCM process. The reactant
gases can be pre-conditioned prior to their introduction into a
catalytic reactor or reactor bed, in a safe and efficient manner.
Such pre-conditioning can include (i) mixing of reactant streams,
such as a methane-containing stream and a stream of an oxidizing
agent (e.g., oxygen) in an OCM reactor or prior to directing the
streams to the OCM reactor, (ii) heating or pre-heating the
methane-containing stream and/or the stream of the oxidizing agent
using, for example, heat from the OCM reactor, or (iii) a
combination of mixing and pre-heating. Such pre-conditioning can
minimize, if not eliminate auto-ignition of methane and the
oxidizing agent. Systems and methods for pre-conditioning reactant
gases are described in, for example, U.S. patent application Ser.
No. 14/553,795, filed Nov. 25, 2014, which is entirely incorporated
herein by reference.
A wide set of competitive reactions can occur simultaneously or
substantially simultaneously with the OCM reaction, including total
combustion of both methane and other partial oxidation products. An
OCM process can yield C.sub.2+ compounds as well as non-C.sub.2+
impurities. The C.sub.2+ compounds can include a variety of
hydrocarbons, such as hydrocarbons with saturated or unsaturated
carbon-carbon bonds. Saturated hydrocarbons can include alkanes,
such as ethane, propane, butane, pentane and hexane. Unsaturated
hydrocarbons may be more suitable for use in downstream non-OCM
processes, such as the manufacture of polymeric materials (e.g.,
polyethylene). Accordingly, it may be preferable to convert at
least some, all or substantially all of the alkanes in the C.sub.2+
compounds to compounds with unsaturated moieties, such as alkenes,
alkynes, alkoxides, ketones, including aromatic variants
thereof.
Once formed, C.sub.2+compounds can be subjected to further
processing to generate desired or otherwise predetermined
chemicals. In some situations, the alkane components of the
C.sub.2+compounds are subjected to cracking in an OCM reactor or a
reactor downstream of the OCM reactor to yield other compounds,
such as alkenes (or olefins). See, e.g., U.S. patent application
Ser. No. 14/553,795, filed Nov. 25, 2014, now published as U.S.
Patent Publication No. 2015/0152025, which is entirely incorporated
herein by reference. FIG. 42 shows an OCM reactor 1501 comprising a
catalyst unit 1502 and a cracking unit 1503. The catalyst unit 1502
can be, for example, a packed bed reactor comprising a
heterogeneous catalyst. The catalyst unit 1502 can be configured to
perform an OCM process using natural gas and O.sub.2 inputted into
the reactor 1501. The cracking unit 1503 is configured to perform
crack alkanes (e.g., ethane) to other types of hydrocarbons, such
as alkenes (e.g., ethylene). The cracking unit 1503 can be
configured to operate adiabatically using heat liberated in the
catalyst unit 1502 in the OCM process, which heat can be conveyed
by way of steam generated in the OCM process, for example.
In some situations, an OCM system generates ethylene that can be
subjected to further processing to generate different hydrocarbons
with the aid of conversion processes (or systems). Such a process
can be part of an ethylene to liquids (ETL) or ethylene, propene,
butene gases to liquids. The ETL process includes OCM olefins
gases, ethylene, propene, butene, or other OCM gaseous products to
produce liquids. OCM-ETL process flow comprising one or more OCM
reactors, separations units, and one or more conversion processes
for generating higher molecular weight hydrocarbons. The conversion
processes can be integrated in a switchable or selectable manner in
which at least a portion or all of the ethylene containing product
can be selectively directed to 1, 2, 3, 4, 5, 6, 7, 8, 9, 10, or
more different process paths to yield as many different hydrocarbon
products. An example OCM and ETL (collectively "OCM-ETL" herein)
process is schematically illustrated in FIG. 1, which shows an OCM
reactor system 100 that includes an OCM reactor train 102 coupled
to an OCM product gas separation train 104. The OCM product gas
separation train 104 can include various separation unit operations
("units"), such as a distillation unit and/or a cryogenic
separation unit. The ethylene rich effluent (shown as arrow 106)
from the separation train 104 is routed to multiple different
ethylene conversion reactor systems and processes 110, e.g.,
ethylene conversion systems 110a-110e, which each produce different
hydrocarbon products, e.g., products 120a-120e. Products 120a-120e
can include, for example, hydrocarbons having between three and
twelve carbon atoms per molecule (C3-C12 hydrocarbons). Such
hydrocarbons may be suitable for use as fuels for various machines,
such as automobiles.
The fluid connection between the OCM reactor system 100 and each of
the different ethylene conversion systems 110a-110e can be
controllable and selective, e.g., with the aid of a valve and
control system, which can apportion the output of the OCM reactor
system 100 to 1, 2, 3, 4, 5, 6, 7, 8, 9, 10, or more different
ethylene conversion systems. The conversions systems 110a-110e can
be ETL or gas to liquids (GTL) reactors. Valve and piping systems
for accomplishing this may take a variety of different forms,
including valves at each piping junction, multiport valves,
multi-valve manifold assemblies, and the like. Other details of the
OCM-ETL process of FIG. 1 are provided in, for example, U.S. patent
application Ser. No. 14/099,614, filed on Dec. 6, 2013, which is
entirely incorporated herein by reference.
As noted, the present disclosure includes processes and systems for
production of various higher hydrocarbons (i.e., C.sub.3+) from
ethylene, and particularly liquid hydrocarbon compositions. In some
aspects, the ethylene is itself derived from methane in a methane
containing feedstock, such as natural gas. Production of ethylene
from methane can be accomplished through a number of different
catalytic pathways, for example in some embodiments, the processes
and systems of the disclosure convert methane to ethylene through
OCM in an OCM reactor system. In some embodiments, the ethylene
produced in the OCM reactor system is charged to one or more
ethylene conversion reactor systems where it can be converted to a
higher hydrocarbon, for example a different higher hydrocarbon in
each of the ethylene conversion reactor systems.
OCM reactions, processes and systems can operate within economic
and reasonable process windows. In some cases, catalysts, processes
and reactor systems have been able to carry out OCM reactions at
temperatures, pressures, selectivities and yields that are
commercially attractive. See, e.g., U.S. patent application Ser.
Nos. 13/115,082, 13/479,767, 13/689,611, 13/739,954, 13/900,898,
13/901,319, 13/936,783, and 14/212,435, the full disclosures of
which are incorporated herein by reference in their entirety for
all purposes.
As used herein, an OCM process or system typically employs one or
more reactor vessels that contain an appropriate OCM catalyst
material, typically in conjunction with additional system
components. A variety of OCM catalysts have been described
previously. See, e.g., U.S. Pat. Nos. 5,712,217, 6,403,523, and
6,576,803, the full disclosures of which are incorporated herein by
reference in their entirety for all purposes. Some catalysts have
been developed that yield conversion and selectivity that enable
economic methane conversion under practical operating conditions.
These are described in, for example, Published U.S. Patent
Application No. 2012-0041246, as well as patent application Ser.
No. 13/479,767, filed May 24, 2012, and Ser. No. 13/689,611, filed
Nov. 29, 2012, the full disclosures of each of which are
incorporated herein by reference in their entirety for all
purposes.
Accordingly, in some embodiments, the disclosure provides a method
of producing a hydrocarbon product, the method comprising: (a)
introducing methane and a source of oxidant into an OCM reactor
system capable of converting methane to ethylene at reactor inlet
temperatures of between about 450.degree. C. and 600.degree. C. and
reactor pressures of between about 15 psig and 125 psig, with
C.sub.2+ selectivity of at least 50%, under conditions for the
conversion of methane to ethylene; (b) converting methane to a
product gas comprising ethylene; (c) introducing at least a portion
of the product gas into an integrated ethylene conversion reaction
systems, the integrated ethylene conversion reaction system being
configured for converting ethylene into a higher hydrocarbon
product; and (d) converting the ethylene into a higher hydrocarbon
product.
In some embodiments, the method is for producing a plurality of
hydrocarbon products. Accordingly, in some embodiments, the
invention provides a method of producing a plurality of hydrocarbon
products, the method comprising: (a) introducing methane and a
source of oxidant into an OCM reactor system capable of converting
methane to ethylene at reactor inlet temperatures of between about
450.degree. C. and 600.degree. C. and reactor pressures of between
about 15 psig and 125 psig, with C.sub.2+ selectivity of at least
50%, under conditions for the conversion of methane to ethylene;
(b) converting methane to a product gas comprising ethylene; (c)
introducing separate portions of the product gas into at least
first and second integrated ethylene conversion reaction systems,
each integrated ethylene conversion reaction system being
configured for converting ethylene into a different higher
hydrocarbon product; and (d) converting the ethylene into different
higher hydrocarbon products. In some embodiments, the integrated
ethylene conversion systems are selected from selective and full
range ethylene conversion systems.
In some embodiments the methods further comprise introducing a
portion of the product gas into at least a third integrated
ethylene conversion system. Some embodiments further comprise
introducing a portion of the product gas into at least first,
second, third and fourth integrated ethylene conversion
systems.
In any of the methods described herein, the integrated ethylene
conversion systems can be selected from linear alpha olefin (LAO)
systems, linear olefin systems, branched olefin systems, saturated
linear hydrocarbon systems, branched hydrocarbon systems, saturated
cyclic hydrocarbon systems, olefinic cyclic hydrocarbon systems,
aromatic hydrocarbon systems, oxygenated hydrocarbon systems,
halogenated hydrocarbon systems, alkylated aromatic systems, and
hydrocarbon polymer systems.
In some embodiments, the integrated ethylene conversion systems can
be selected from LAO systems that produce one or more of 1-butene,
1-hexene, 1-octene and 1-decene. For example, in certain
embodiments at least one of the LAO systems is configured for
performing a selective LAO process.
In some embodiments, at least one of the integrated ethylene
conversion systems comprises a full range ethylene oligomerization
system configured for producing higher hydrocarbons in the range of
C.sub.4 to C.sub.30.
In some embodiments, the OCM reactor system comprises nanowire OCM
catalyst material. In some embodiments, the product gas comprises
less than 5 mol % of ethylene. For example, in certain embodiments,
the product gas comprises less than 3 mol % of ethylene. In some
embodiments, the product gas can further comprise one or more gases
selected from CO.sub.2, CO, H.sub.2, H.sub.2O, C.sub.2H.sub.6,
CH.sub.4 and C.sub.3+ hydrocarbons.
In some embodiments, the method further comprises enriching the
product gas for ethylene prior to introducing the separate portions
of the product gas into the at least first and second integrated
ethylene conversion reaction systems.
In some embodiments, the method further comprises introducing an
effluent gas from the first or second integrated ethylene
conversion reaction systems into the OCM reactor system. For
example, in some of these embodiments the method further comprises
converting methane present in the effluent gas to ethylene and
charging the ethylene to one or more of the integrated ethylene
conversion systems.
In various embodiments, the disclosure is directed to a method of
producing a plurality of hydrocarbon products, the method
comprising: (a) introducing methane and a source of oxidant into an
OCM reactor system capable of converting methane to ethylene at
reactor inlet temperatures of between about 450.degree. C. and
600.degree. C. and reactor pressures of between about 15 psig and
125 psig, with C.sub.2+ selectivity of at least 50%, under
conditions for the conversion of methane to ethylene; (b)
recovering ethylene from the OCM reactor system; and (c)
introducing separate portions of the ethylene recovered from the
OCM reactor system into at least two integrated, but discrete and
different catalytic ethylene conversion reaction systems for
converting ethylene into at least two different higher hydrocarbon
products.
In some embodiments, the at least two ethylene conversion systems
are selected from selective and full range ethylene conversion
systems. In some embodiments, the at least two ethylene conversion
systems comprise at least three ethylene conversion systems. For
example, in some embodiments the at least two ethylene conversion
systems comprise at least four ethylene conversion systems.
In yet more embodiments, the at least two ethylene conversion
systems are selected from linear alpha olefin (LAO) systems, linear
olefin systems, branched olefin systems, saturated linear
hydrocarbon systems, branched hydrocarbon systems, saturated cyclic
hydrocarbon systems, olefinic cyclic hydrocarbon systems, aromatic
hydrocarbon systems, oxygenated hydrocarbon systems, halogenated
hydrocarbon systems, alkylated aromatic systems, and hydrocarbon
polymer systems.
In some cases, the at least two ethylene conversion systems are
selected from LAO systems that produce one or more of 1-butene,
1-hexene, 1-octene and 1-decene. For example, in some embodiments
at least one of the at least two LAO processes comprises a
selective LAO process, and in other exemplary embodiments at least
one of the at least two ethylene conversion systems comprises a
full range ethylene oligomerization system for producing higher
hydrocarbons in the range of C.sub.4 to C.sub.30. In some
instances, the OCM reactor system comprises nanowire OCM catalyst
material.
In some embodiments, the disclosure provides a method of producing
a plurality of liquid hydrocarbon products, comprising: (a)
converting methane to a product gas comprising ethylene using a
catalytic reactor process; and (b) contacting separate portions of
the product gas with at least two discrete catalytic reaction
systems selected from linear alpha olefin (LAO) systems, linear
olefin systems, branched olefin systems, saturated linear
hydrocarbon systems, branched hydrocarbon systems, saturated cyclic
hydrocarbon systems, olefinic cyclic hydrocarbon systems, aromatic
hydrocarbon systems, oxygenated hydrocarbon systems, halogenated
hydrocarbon systems, alkylated aromatic systems, and hydrocarbon
polymer systems.
In some cases, a method of producing a plurality of liquid
hydrocarbon products is provided. The method comprises: (a)
converting methane to ethylene using a catalytic reactor process;
(b) recovering ethylene from the catalytic reactor process; and (c)
contacting separate portions of the ethylene recovered from the OCM
reactor system with at least two discrete catalytic reaction
systems selected from linear alpha olefin (LAO) systems, linear
olefin systems, branched olefin systems, saturated linear
hydrocarbon systems, branched hydrocarbon systems, saturated cyclic
hydrocarbon systems, olefinic cyclic hydrocarbon systems, aromatic
hydrocarbon systems, oxygenated hydrocarbon systems, halogenated
hydrocarbon systems, alkylated aromatic systems, and hydrocarbon
polymer systems.
Some embodiments of the present disclosure are directed to a
processing system for preparation of C.sub.2+ hydrocarbon products
from methane. For example, in some embodiments the invention
provides a processing system comprising: (a) an OCM reactor system
comprising an OCM catalyst, the OCM reactor system being fluidly
connected at an input, to a source of methane and a source of
oxidant; (b) an integrated ethylene conversion reactor system, the
ethylene reactor system being configured to convert ethylene to a
higher hydrocarbon; and (c) a selective coupling between the OCM
reactor system and the ethylene reactor system, the selective
coupling configured to selectively direct a portion or all of the
product gas to the ethylene conversion reactor system.
In some instances, the disclosure provides a processing system
comprising: (a) an OCM reactor system comprising an OCM catalyst,
the OCM reactor system being fluidly connected at an input, to a
source of methane and a source of oxidant; (b) at least first and
second catalytic ethylene conversion reactor systems, the first
catalytic ethylene reactor system being configured to convert
ethylene to a first higher hydrocarbon, and the second catalytic
ethylene reactor system being configured to convert ethylene to a
second higher hydrocarbon different from the first higher
hydrocarbon; and (c) a selective coupling between the OCM reactor
system and the first and second catalytic ethylene reactor systems
configured to selectively direct a portion or all of the product
gas to each of the first and second catalytic ethylene reactor
systems.
In some embodiments, the ethylene conversion systems are selected
from linear alpha olefin (LAO) systems, linear olefin systems,
branched olefin systems, saturated linear hydrocarbon systems,
branched hydrocarbon systems, saturated cyclic hydrocarbon systems,
olefinic cyclic hydrocarbon systems, aromatic hydrocarbon systems,
oxygenated hydrocarbon systems, halogenated hydrocarbon systems,
alkylated aromatic systems, ethylene copolymerization systems, and
hydrocarbon polymer systems.
In some instances, the OCM catalyst comprises a nanowire catalyst.
In some embodiments, the system further comprises an ethylene
recovery system fluidly coupled between the OCM reactor system and
the at least first and second catalytic ethylene conversion reactor
systems, the ethylene recovery system configured for enriching the
product gas for ethylene.
In some cases, the disclosure is directed to a processing system,
the processing system comprising: (a) an OCM reactor system
comprising an OCM catalyst, the OCM reactor system being fluidly
connected at an input, to a source of methane and a source of
oxidant; (b) an ethylene recovery system fluidly coupled to the OCM
reactor system at an outlet, for recovering ethylene from an OCM
product gas; (c) at least first and second catalytic ethylene
conversion reactor systems, the first catalytic ethylene reactor
system being configured to convert ethylene to a first higher
hydrocarbon composition, and the second catalytic ethylene reactor
system being configured to convert ethylene to a second higher
hydrocarbon composition different from the first higher hydrocarbon
composition; and (d) a selective coupling between the outlet of the
ethylene recovery system and the first and second catalytic
ethylene reactor systems to selectively direct a portion or all of
the ethylene recovered from the OCM product gas to each of the
first and second catalytic ethylene reactor systems.
In some cases, two or more of the at least two ethylene conversion
systems are selected from linear alpha olefin (LAO) systems, linear
olefin systems, branched olefin systems, saturated linear
hydrocarbon systems, branched hydrocarbon systems, saturated cyclic
hydrocarbon systems, olefinic cyclic hydrocarbon systems, aromatic
hydrocarbon systems, oxygenated hydrocarbon systems, halogenated
hydrocarbon systems, alkylated aromatic systems, ethylene
copolymerization systems, and hydrocarbon polymer systems. In other
embodiments, the OCM catalyst comprises a nanowire catalyst.
In some embodiments, the catalyst systems used in any of the above
described OCM reaction comprise nanowire catalysts. Such nanowire
catalysts can include substantially straight nanowires or nanowires
having a curved, twisted or bent morphology. The actual lengths of
the nanowire catalysts may vary. For example in some embodiments,
the nanowires have an actual length of between 100 nm and 100
.mu.m. In other embodiments, the nanowires have an actual length of
between 100 nm and 10 .mu.m. In other embodiments, the nanowires
have an actual length of between 200 nm and 10 .mu.m. In other
embodiments, the nanowires have an actual length of between 500 nm
and 5 .mu.m. In other embodiments, the actual length is greater
than 5 .mu.m. In other embodiments, the nanowires have an actual
length of between 800 nm and 1000 nm. In other further embodiments,
the nanowires have an actual length of 900 nm. As noted below, the
actual length of the nanowires may be determined by TEM, for
example, in bright field mode at 5 keV.
The diameter of the nanowires may be different at different points
along the nanowire backbone. However, the nanowires comprise a mode
diameter (i.e., the most frequently occurring diameter). As used
herein, the diameter of a nanowire refers to the mode diameter. In
some embodiments, the nanowires have a diameter of between 1 nm and
10 .mu.m, between 1 nm and 1 .mu.m, between 1 nm and 500 nm,
between 1 nm and 100 nm, between 7 nm and 100 nm, between 7 nm and
50 nm, between 7 nm and 25 nm, or between 7 nm and 15 nm. On other
embodiments, the diameter is greater than 500 nm. As noted below,
the diameter of the nanowires may be determined by TEM, for
example, in bright field mode at 5 keV.
The nanowire catalysts may have different aspect ratios. In some
embodiments, the nanowires have an aspect ratio of greater than
10:1. In other embodiments, the nanowires have an aspect ratio
greater than 20:1. In other embodiments, the nanowires have an
aspect ratio greater than 50:1. In other embodiments, the nanowires
have an aspect ratio greater than 100:1.
In some embodiments, the nanowires comprise a solid core while in
other embodiments, the nanowires comprise a hollow core. In
general, the morphology of a nanowire (including length, diameter,
and other parameters) can be determined by transmission electron
microscopy (TEM). Transmission electron microscopy (TEM) is a
technique whereby a beam of electrons is transmitted through an
ultra-thin specimen, interacting with the specimen as it passes
through. An image is formed from the interaction of the electrons
transmitted through the specimen. The image is magnified and
focused onto an imaging device, such as a fluorescent screen, on a
layer of photographic film or detected by a sensor such as a CCD
camera.
In some embodiments, the nanowire catalysts comprise one or
multiple crystal domains, e.g., monocrystalline or polycrystalline,
respectively. In some other embodiments, the average crystal domain
of the nanowires is less than 100 nm, less than 50 nm, less than 30
nm, less than 20 nm, less than 10 nm, less than 5 nm, or less than
2 nm. Crystal structure, composition, and phase, including the
crystal domain size of the nanowires, can be determined by XRD.
Typically, the nanowire catalytic material comprises a plurality of
nanowires. In certain embodiments, the plurality of nanowires form
a mesh of randomly distributed and, to various degrees,
interconnected nanowires, that presents a porous matrix.
The total surface area per gram of a nanowire or plurality of
nanowires may have an effect on the catalytic performance. Pore
size distribution may affect the nanowires catalytic performance as
well. Surface area and pore size distribution of the nanowires or
plurality of nanowires can be determined by BET (Brunauer, Emmett,
Teller) measurements. BET techniques utilize nitrogen adsorption at
various temperatures and partial pressures to determine the surface
area and pore sizes of catalysts. There are BET techniques for
determining surface area and pore size distribution currently
available. In some embodiments the nanowires have a surface area of
between 0.0001 and 3000 m.sup.2/g, between 0.0001 and 2000
m.sup.2/g, between 0.0001 and 1000 m.sup.2/g, between 0.0001 and
500 m.sup.2/g, between 0.0001 and 100 m.sup.2/g, between 0.0001 and
50 m.sup.2/g, between 0.0001 and 20 m.sup.2/g, between 0.0001 and
10 m.sup.2/g or between 0.0001 and 5 m.sup.2/g. In some embodiments
the nanowires have a surface area of between 0.001 and 3000
m.sup.2/g, between 0.001 and 2000 m.sup.2/g, between 0.001 and 1000
m.sup.2/g, between 0.001 and 500 m.sup.2/g, between 0.001 and 100
m.sup.2/g, between 0.001 and 50 m.sup.2/g, between 0.001 and 20
m.sup.2/g, between 0.001 and 10 m.sup.2/g or between 0.001 and 5
m.sup.2/g. In some other embodiments the nanowires have a surface
area of between 2000 and 3000 m.sup.2/g, between 1000 and 2000
m.sup.2/g, between 500 and 1000 m.sup.2/g, between 100 and 500
m.sup.2/g, between 10 and 100 m.sup.2/g, between 5 and 50
m.sup.2/g, between 2 and 20 m.sup.2/g or between 0.0001 and 10
m.sup.2/g. In other embodiments, the nanowires have a surface area
of greater than about 2000 m.sup.2/g, greater than about 1000
m.sup.2/g, greater than about 500 m.sup.2/g, greater than about 100
m.sup.2/g, greater than about 50 m.sup.2/g, greater than about 20
m.sup.2/g, greater than about 10 m.sup.2/g, greater than about 5
m.sup.2/g, greater than about 1 m.sup.2/g, greater than about
0.0001 m.sup.2/g.
The nanowire catalysts and catalyst compositions used in
conjunction with the processes and systems of some embodiments of
the invention may have any number of compositions and/or
morphologies. These nanowire catalysts may be inorganic and either
polycrystalline or mono-crystalline. In some other embodiments, the
nanowires are inorganic and polycrystalline. In certain examples,
the nanowire catalysts comprise one or more elements from any of
Groups 1 through 7, lanthanides, actinides or combinations thereof.
Thus in certain aspects, the catalysts comprise an inorganic
catalytic polycrystalline nanowire, the nanowire having a ratio of
effective length to actual length of less than one and an aspect
ratio of greater than ten as measured by TEM in bright field mode
at 5 keV, wherein the nanowire comprises one or more elements from
any of Groups 1 through 7, lanthanides, actinides or combinations
thereof.
In still other cases, the nanowire catalysts comprise one or more
metal elements from any of Groups 1-7, lanthanides, actinides or
combinations thereof, for example, the nanowires may be
mono-metallic, bi-metallic, tri-metallic, etc. (i.e., contain one,
two, three, etc. metal elements), where the metal elements may be
present in the nanowires in elemental or oxidized form, or in the
form of a compound comprising a metal element. The metal element or
compound comprising the metal element may be in the form of oxides,
hydroxides, oxyhydroxides, salts, hydrated oxides, carbonates,
oxy-carbonates, sulfates, phosphates, acetates, oxalates and the
like. The metal element or compound comprising the metal element
may also be in the form of any of a number of different polymorphs
or crystal structures.
In some examples, metal oxides may be hygroscopic and may change
forms once exposed to air, may absorb carbon dioxide, may be
subjected to incomplete calcination or any combination thereof.
Accordingly, although the nanowires are often referred to as metal
oxides, in some embodiments the nanowires also comprise hydrated
oxides, oxyhydroxides, hydroxides, oxycarbonates (or oxide
carbonates), carbonates or combinations thereof.
In some cases, the nanowires comprise one or more metal elements
from Group 1, Group 2, Group 3, Group 4, Group 5, Group 6, Group 7,
lanthanides, and/or actinides, or combinations of these, as well as
oxides of these metals. In other cases, the nanowires comprise
hydroxides, sulfates, carbonates, oxide carbonates, acetates,
oxalates, phosphates (including hydrogen phosphates and
dihydrogenphosphates), oxy-carbonates, oxyhalides, hydroxyhalides,
oxyhydroxides, oxysulfates, mixed oxides or combinations thereof of
one or more metal elements from any of Groups 1-7, lanthanides,
actinides or combinations thereof. Examples of such nanowire
materials include, but are not limited to nanowires comprising,
e.g., Li.sub.2CO.sub.3, LiOH, Li.sub.2O, Li.sub.2C.sub.2O.sub.4,
Li.sub.2SO.sub.4, Na.sub.2CO.sub.3, NaOH, Na.sub.2O,
Na.sub.2C.sub.2O.sub.4, Na.sub.2SO.sub.4, K.sub.2CO.sub.3, KOH,
K.sub.2O, K.sub.2C.sub.2O.sub.4, K.sub.2SO.sub.4, Cs.sub.2CO.sub.3,
CsOH, Cs.sub.2O, CsC.sub.2O.sub.4, CsSO.sub.4, Be(OH).sub.2,
BeCO.sub.3, BeO, BeC.sub.2O.sub.4. BeSO.sub.4, Mg(OH).sub.2,
MgCO.sub.3, MgO, MgC.sub.2O.sub.4. MgSO.sub.4, Ca(OH).sub.2, CaO,
CaCO.sub.3, CaC.sub.2O.sub.4, CaSO.sub.4, Y.sub.2O.sub.3,
Y.sub.2(CO.sub.3).sub.3, Y(OH).sub.3, Y.sub.2(C.sub.2O4).sub.3,
Y.sub.2(SO.sub.4).sub.3, Zr(OH).sub.4, ZrO(OH).sub.2, ZrO2,
Zr(C.sub.2O.sub.4).sub.2, Zr(SO.sub.4).sub.2, Ti(OH).sub.4,
TiO(OH).sub.2, TiO.sub.2, Ti(C.sub.2O.sub.4).sub.2,
Ti(SO.sub.4).sub.2, BaO, Ba(OH).sub.2, BaCO.sub.3,
BaC.sub.2O.sub.4, BaSO.sub.4, La(OH).sub.3, La.sub.2O.sub.3,
La.sub.2(C.sub.2O.sub.4).sub.3, La.sub.2(SO.sub.4).sub.3,
La.sub.2(CO.sub.3).sub.3, Ce(OH).sub.4, CeO.sub.2, Ce.sub.2O.sub.3,
Ce(C.sub.2O.sub.4).sub.2, Ce(SO.sub.4).sub.2, Ce(CO.sub.3).sub.2,
ThO.sub.2, Th(OH).sub.4, Th(C.sub.2O.sub.4).sub.2,
Th(SO.sub.4).sub.2, Th(CO.sub.3).sub.2, Sr(OH).sub.2, SrCO.sub.3,
SrO, SrC.sub.2O.sub.4, SrSO.sub.4, Sm.sub.2O.sub.3, Sm(OH).sub.3,
Sm.sub.2(CO.sub.3).sub.3, Sm.sub.2(C.sub.2O.sub.4).sub.3,
Sm.sub.2(SO.sub.4).sub.3, LiCa.sub.2Bi.sub.3O.sub.4Cl.sub.6,
NaMnO.sub.4, Na.sub.2WO.sub.4, NaMn/WO.sub.4, CoWO.sub.4,
CuWO.sub.4, K/SrCoO.sub.3, K/Na/SrCoO.sub.3, Na/SrCoO.sub.3,
Li/SrCoO.sub.3, SrCoO.sub.3, Mg.sub.6MnO.sub.8, LiMn.sub.2O.sub.4,
Li/Mg.sub.6MnO.sub.8, Na.sub.10Mn/W.sub.5O.sub.17,
Mg.sub.3Mn.sub.3B.sub.2O.sub.10, Mg.sub.3(BO.sub.3).sub.2,
molybdenum oxides, molybdenum hydroxides, molybdenum oxalates,
molybdenum sulfates, Mn.sub.2O.sub.3, Mn.sub.3O.sub.4, manganese
oxides, manganese hydroxides, manganese oxalates, manganese
sulfates, manganese tungstates, manganese carbonates, vanadium
oxides, vanadium hydroxides, vanadium oxalates, vanadium sulfates,
tungsten oxides, tungsten hydroxides, tungsten oxalates, tungsten
sulfates, neodymium oxides, neodymium hydroxides, neodymium
carbonates, neodymium oxalates, neodymium sulfates, europium
oxides, europium hydroxides, europium carbonates, europium
oxalates, europium sulfates, praseodymium oxides, praseodymium
hydroxides, praseodymium carbonates, praseodymium oxalates,
praseodymium sulfates, rhenium oxides, rhenium hydroxides, rhenium
oxalates, rhenium sulfates, chromium oxides, chromium hydroxides,
chromium oxalates, chromium sulfates, potassium molybdenum
oxides/silicon oxide or combinations thereof.
Still other examples of these nanowire materials include, but are
not limited to, nanowires comprising, e.g., Li.sub.2O, Na.sub.2O,
K.sub.2O, Cs.sub.2O, BeO MgO, CaO, ZrO(OH).sub.2, ZrO.sub.2,
TiO.sub.2, TiO(OH).sub.2, BaO, Y.sub.2O.sub.3, La.sub.2O.sub.3,
CeO.sub.2, Ce.sub.2O3, ThO.sub.2, SrO, Sm.sub.2O.sub.3,
Nd.sub.2O.sub.3, Eu.sub.2O.sub.3, Pr.sub.2O.sub.3,
LiCa.sub.2Bi.sub.3O.sub.4C.sub.16, NaMnO.sub.4, Na.sub.2WO.sub.4,
Na/Mn/WO.sub.4, Na/MnWO.sub.4, Mn/WO.sub.4, K/SrCoO.sub.3,
K/Na/SrCoO.sub.3, K/SrCoO.sub.3, Na/SrCoO.sub.3, Li/SrCoO.sub.3,
SrCoO.sub.3, Mg.sub.6MnO.sub.8, Na/B/Mg.sub.6MnO.sub.8,
Li/B/Mg.sub.6MnO.sub.8, Zr.sub.2Mo.sub.2O.sub.8, molybdenum oxides,
Mn.sub.2O.sub.3, Mn.sub.3O.sub.4, manganese oxides, vanadium
oxides, tungsten oxides, neodymium oxides, rhenium oxides, chromium
oxides, or combinations thereof. A variety of different nanowire
compositions have been described in, e.g., Published U.S. Patent
Application No. 2012-0041246 and U.S. patent application Ser. No.
13/689,611, filed Nov. 29, 2012 (the full disclosures of which are
incorporated herein in their entirety for all purposes), and are
envisioned for use in conjunction with the present invention.
Products produced from these catalytic reactions typically include
CO, CO.sub.2, H.sub.2O, C.sub.2+ hydrocarbons, such as ethylene,
ethane, and larger alkanes and alkenes, such as propane and
propylene. In some embodiments, the OCM reactor systems operate to
convert methane into desired higher hydrocarbon products (ethane,
ethylene, propane, propylene, butanes, pentanes, etc.),
collectively referred to as C.sub.2+ compounds, with high yield. In
particular, the progress of the OCM reaction is generally discussed
in terms of methane conversion, C.sub.2+ selectivity, and C.sub.2+
yield. As used herein, methane conversion generally refers to the
percentage or fraction of methane introduced into the reaction that
is converted to a product other than methane. C.sub.2+ selectivity
generally refers to the percentage of all non-methane, carbon
containing products of the OCM reaction that are the desired
C.sub.2+ products, e.g., ethane, ethylene, propane, propylene, etc.
Although primarily stated as C.sub.2+ selectivity, it will be
appreciated that selectivity may be stated in terms of any of the
desired products, e.g., just C2, or just C2 and C3. Finally,
C.sub.2+ yield generally refers to the amount of carbon that is
incorporated into a C.sub.2+ product as a percentage of the amount
of carbon introduced into a reactor in the form of methane. This
may generally be calculated as the product of the conversion and
the selectivity divided by the number of carbon atoms in the
desired product. C.sub.2+ yield is typically additive of the yield
of the different C.sub.2+ components included in the C.sub.2+
components identified, e.g., ethane yield+ethylene yield+propane
yield+propylene yield etc.).
Exemplary OCM processes and systems typically provide a methane
conversion of at least 10% per process pass in a single integrated
reactor system (e.g., single isothermal reactor system or
integrated multistage adiabatic reactor system), with a C.sub.2+
selectivity of at least 50%, but at reactor inlet temperatures of
between 400 and 600.degree. C. and at reactor inlet pressures of
between about 15 psig and about 150 psig. Thus, the catalysts
employed within these reactor systems are capable of providing the
described conversion and selectivity under the described reactor
conditions of temperature and pressure. In the context of some OCM
catalysts and system embodiments, it will be appreciated that the
reactor inlet or feed temperatures typically substantially
correspond to the minimum "light-off" or reaction initiation
temperature for the catalyst or system. Restated, the feed gases
are contacted with the catalyst at a temperature at which the OCM
reaction is able to be initiated upon introduction to the reactor.
Because the OCM reaction is exothermic, once light-off is achieved,
the heat of the reaction can be expected to maintain the reaction
at suitable catalytic temperatures, and even generate excess
heat.
In some aspects, the OCM reactors and reactor systems, when
carrying out the OCM reaction, operate at pressures of between
about 15 psig and about 125 psig at the above described
temperatures, while providing the conversion and selectivity
described herein, and in even more embodiments, at pressures less
than 100 psig, e.g., between about 15 psig and about 100 psig.
Examples of particularly useful catalyst materials are described
in, for example, Published U.S. Patent Application No.
2012-0041246, as well as patent application Ser. No. 13/479,767,
filed May 24, 2012, and Ser. No. 13/689,611, filed Nov. 29, 2012,
which are incorporated herein by reference in their entirety for
all purposes. In some embodiments, the catalysts comprise bulk
catalyst materials, e.g., having relatively undefined morphology
or, in certain embodiments, the catalyst material comprises, at
least in part, nanowire containing catalytic materials. In either
form, the catalysts used in accordance with the present invention
may be employed under the full range of reaction conditions
described above, or in any narrower described range of conditions.
Similarly, the catalyst materials may be provided in a range of
different larger scale forms and formulations, e.g., as mixtures of
materials having different catalytic activities, mixtures of
catalysts and relatively inert or diluent materials, incorporated
into extrudates, pellets, or monolithic forms, or the like. Ranges
of exemplary catalyst forms and formulations are described in, for
example, U.S. patent application Ser. No. 13/901,319 and U.S.
Provisional Patent Application No. 62/051,779, the full disclosures
of which are incorporated herein by reference in their entireties
for all purposes.
The reactor vessels used for carrying out the OCM reaction in the
OCM reactor systems of the invention may include one or more
discrete reactor vessels each containing OCM catalyst material,
fluidly coupled to a methane source and a source of oxidant as
further discussed elsewhere herein. Feed gas containing methane
(e.g., natural gas) is contacted with the catalyst material under
conditions suitable for initiation and progression of the reaction
within the reactor to catalyze the conversion of methane to
ethylene and other products.
For example, in some embodiments the OCM reactor system comprises
one or more staged reactor vessels operating under isothermal or
adiabatic conditions, for carrying out OCM reactions. For adiabatic
reactor systems, the reactor systems may include one, two, three,
four, five or more staged reactor vessels arranged in series, which
are fluidly connected such that the effluent or "product gas" of
one reactor is directed, at least in part, to the inlet of a
subsequent reactor. Such staged serial reactors provide higher
yield for the overall process, by allowing catalytic conversion of
previously unreacted methane. These adiabatic reactors are
generally characterized by the lack of an integrated thermal
control system used to maintain little or no temperature gradient
across the reactor. With no integrated temperature control system,
the exothermic nature of the OCM reaction results in a temperature
gradient across the reactor indicative of the progress of the
reaction, where the inlet temperature can range from about
450.degree. C. to about 600.degree. C., while the outlet
temperature ranges from about 700.degree. C. to about 900.degree.
C. Typically, such temperature gradients can range from about
100.degree. C. to about 450.degree. C. By staging adiabatic
reactors, with interstage cooling systems, one can step through a
more complete catalytic reaction without generating extreme
temperatures, e.g., in excess of 900.degree. C.
In operation of certain embodiments, methane-containing feed gas is
introduced into the inlet side of a reactor vessel, e.g., the first
reactor in a staged reactor system. Within this reactor, the
methane is converted into C.sub.2+ hydrocarbons, as well as other
products, as discussed above. At least a portion of the product gas
stream is then cooled to an appropriate temperature and introduced
into a subsequent reactor stage for continuation of the catalytic
reaction. In particular, the effluent from a preceding reactor,
which in some cases may include unreacted methane, can provide at
least a portion of the methane source for a subsequent reactor. An
oxidant source and a methane source, separate from the unreacted
methane from the first reactor stage, are also typically coupled to
the inlet of each subsequent reactor.
In some aspects, the reactor systems include one or more
`isothermal` reactors, that maintain a relatively low temperature
gradient across the overall reactor bed, e.g., between the inlet
gas and outlet or product gas, through the inclusion of integrated
temperature control elements, such as coolant systems that contact
heat exchange surfaces on the reactor to remove excess heat, and
maintain a flat or insignificant temperature gradient between the
inlet and outlet of the reactor. Typically, such reactors utilize
molten salt or other coolant systems that operate at temperatures
below 593.degree. C. As with adiabatic systems, isothermal reactor
systems may include one, two, three or more reactors that may be
configured in serial or parallel orientation. Reactor systems for
carrying out these catalytic reactions are also described in U.S.
patent application Ser. No. 13/900,898, the full disclosure of
which is incorporated herein by reference in its entirety for all
purposes.
The OCM reactor systems used in certain embodiments of the present
invention also typically include thermal control systems that are
configured to maintain a desired thermal or temperature profile
across the overall reactor system, or individual reactor vessels.
In the context of adiabatic reactor systems, it will be appreciated
that the thermal control systems include, for example, heat
exchangers disposed upstream, downstream or between serial reactors
within the overall system in order to maintain the desired
temperature profile across the one or more reactors. In the context
of reactors carrying out exothermic reactions, like OCM, such
thermal control systems also optionally include control systems for
modulating flow of reactants, e.g., methane containing feed gases
and oxidant, into the reactor vessels in response to temperature
information feedback, in order to modulate the reactions to achieve
the thermal profiles of the reactors within the desired temperature
ranges. These systems are also described in U.S. patent application
Ser. No. 13/900,898, previously incorporated herein by
reference.
For isothermal reactors, such thermal control systems include the
foregoing, as well as integrated heat exchange components, such as
integrated heat exchangers built into the reactors, such as
tube/shell reactor/heat exchangers, where a void space is provided
surrounding a reactor vessel or through which one or more reactor
vessels or tubes pass. A heat exchange medium is then passed
through the void to remove heat from the individual reactor tubes.
The heat exchange medium is then routed to an external heat
exchanger to cool the medium prior to recirculation into the
reactor.
Following the OCM process, ethylene optionally may be recovered
from the OCM product gas using an ethylene recovery process that
separates ethylene present in the product gas from other
components, such as residual, i.e., unreacted methane, ethane, and
higher hydrocarbons, such as propane, butanes, pentanes and the
like. Alternatively, the OCM product gas is used in subsequent
reactions, as described below, without further purification or
separation of the ethylene. In various other embodiments, the OCM
product gas is enriched for ethylene before being used in
subsequent reactions. In this respect, "enriched" includes, but is
not limited to, operations which increases the overall mol % of
ethylene in the product gas.
In accordance with the present disclosure, ethylene derived from
methane, e.g., using the OCM processes and systems, is further
processed into higher hydrocarbon compositions, and particularly
liquid hydrocarbon compositions. For ease of discussion, reference
to OCM processes and systems, when referring to their inclusion in
an overall process flow, from methane to higher hydrocarbon
compositions, also optionally includes intermediate process steps
involved in purification of ethylene from an OCM product gas, e.g.,
recycling of product gases through the OCM reactor system,
separations of methane and higher hydrocarbons, e.g., NGLs and
other C.sub.2+ compounds, from the OCM product gas, and the like.
Examples of such intermediate processes include, for example,
cryogenic or lean oil separation systems, temperature swing
adsorption (TSA), pressure swing adsorption (PSA), and membrane
separations, for separation of different hydrocarbon and other
components from ethylene, e.g., CO, CO.sub.2, water, nitrogen,
residual methane, ethane, propane, and other higher hydrocarbon
compounds, potentially present in the OCM product gas. Examples of
such systems are described in, e.g., U.S. patent application Ser.
Nos. 13/739,954, 13/936,783, and 13/936,870, the full disclosures
of which are incorporated herein by reference in their entirety for
all purposes.
FIG. 10 schematically illustrates an exemplary OCM system with
integrated separations system component or components. In
particular, shown in FIG. 10 is an exemplary process flow diagram
depicting a process 1000 for methane based C.sub.2 production, in a
product gas from an OCM reactor or reactors 1002, and separation
process 1004, that includes a first separator 1006 providing the
C.sub.2-rich effluent 1052 and a methane/nitrogen-rich effluent
1074. In the embodiment illustrated in FIG. 10, the OCM product gas
from the OCM reactor(s) 1002 is compressed through compressor 1026.
The temperature of the compressed OCM product gas 1050 is reduced
using one or more heat exchangers 1010. The temperature of the
compressed OCM product gas 1050 may be reduced through the use of
an external provided cooling media, introduction of or thermal
exchange with a cool process stream, or combinations of these.
Reducing the temperature of the OCM product gas 1050 will typically
condense at least a portion of the higher boiling point components
in the compressed OCM product gas 1050, including at least a
portion of the C.sub.2 and heavier hydrocarbon components present
in the compressed OCM product gas 1050.
At least a portion of the condensed high boiling point components
can be separated from the compressed OCM product gas 1050 using one
or more liquid gas separators, such as knockout drums 1012 to
provide an OCM product gas condensate 1054 and a compressed OCM
product gas 1056. The OCM product gas condensate 1054 is introduced
to the first separator 1006 and at least a portion 1058 of the
compressed OCM product gas 1056 can be introduced to one or more
turboexpanders 1014. The isentropic expansion of the compressed OCM
product gas 1058 within the turboexpanders 1014 can produce shaft
work useful for driving one or more compressors or other devices in
the separation unit 1004. The isentropic expansion of the
compressed OCM product gas 1058 with the turboexpanders reduces the
temperature of the compressed OCM product gas 1060 that exits from
the one or more turboexpanders. The compressed OCM product gas 1060
from the one or more turboexpanders 1014 is introduced to the first
separator 1006.
The first separator 1006 can be any system, device or combination
of systems and devices suitable for promoting the separation of
C.sub.2 and heavier hydrocarbons from a gas stream that includes
methane and nitrogen. For example, cryogenic distillation at a
relatively high temperature may be used to promote separation of
the C.sub.2 and heavier hydrocarbons from the methane and nitrogen
components in the gas stream. The C.sub.2-rich effluent 1052 is
withdrawn from the first separator 1006 and a mixed nitrogen and
methane containing gas mixture 1074 is also withdrawn from the
first separator 1054. The nitrogen content of the nitrogen/methane
containing gas mixture 1074 withdrawn from the first separator 1006
can be about 95 mol % or less; about 85 mol % or less; about 75 mol
% or less; about 55 mol % or less; about 30 mol % or less. The
balance of the nitrogen/methane gas mixture 1054 comprises
principally methane with small quantities of hydrogen, carbon
monoxide, and inert gases such as argon. The nitrogen/methane rich
gas 1074 is then further cooled using heat exchanger(s) 1022, and
the cooled nitrogen/methane containing gas 1076 is then introduced
into second separator 1008, described in more detail, below.
In at least some embodiments, the first separator functions as a
"demethanizer" based upon its ability to separate methane from the
C.sub.2 and heavier hydrocarbon components. An exemplary first
separator 1006 includes a vertical distillation column operating at
below ambient temperature and above ambient pressure. In
particular, the operating temperature and pressure within the first
separator 1006 can be established to improve the recovery of the
desired C.sub.2 hydrocarbons in the C.sub.2-rich effluent 1052. In
exemplary embodiments, the first separator 1006 can have an
overhead operating temperature of from about -260.degree. F.
(-162.degree. C.) to about -180.degree. F. (-118.degree. C.); about
-250.degree. F. (-157.degree. C.) to about -190.degree. F.
(-123.degree. C.); about -240.degree. F. (-151.degree. C.) to about
-200.degree. F. (-129.degree. C.; or even from about -235.degree.
F. (-148.degree. C.) to about -210.degree. F. (-134.degree. C.) and
a bottom operating temperature of from about -150.degree. F.
(-101.degree. C.) to about -50.degree. F. (-46.degree. C.); about
-135.degree. F. (-93.degree. C.) to about -60.degree. F.
(-51.degree. C.); from about -115.degree. F. (-82.degree. C.) to
about -70.degree. F. (-57.degree. C.); or about -100.degree. F.
(-73.degree. C.) to about -80.degree. F. (-62.degree. C.). In an
exemplary aspect, the first separator 1006 may operate at pressures
of from about 30 psig (205 kPa) to about 130 psig (900 kPa); about
40 psig (275 kPa) to about 115 psig (790 kPa); about 50 psig (345
kPa) to about 95 psig (655 kPa); or about 60 psig (415 kPa) to
about 80 psig (550 kPa).
The temperature of at least a portion of the C.sub.2-rich effluent
1052 from the first separator 1006 can be increased in one or more
heat exchangers 1016, again using an externally supplied heat
transfer medium, introduction of, or thermal contact, with a warmer
process flow stream, or a combination of these, or other heating
systems. The one or more heat exchanger devices 1016 may include
any type of heat exchange device or system, including but not
limited to one or more plate and frame, shell and tube or similar
heat exchanger system. After exiting the one or more heat
exchangers 1016, the heated C.sub.2-rich effluent 1052 may be at
temperatures of 50.degree. F. (10.degree. C.) or less; 25.degree.
F. (-4.degree. C.) or less; about 0.degree. F. (-18.degree. C.) or
less; about -25.degree. F. (-32.degree. C.) or less; or about
-50.degree. F. (-46.degree. C.) or less. Furthermore, the pressure
may be about 130 psig (900 kPa) or less; about 115 psig (790 kPa or
less; about 100 psig (690 kPa) or less; or about 80 psig (550 kPa)
or less.
In some embodiments, a portion 1062 of the OCM product gas 1056
removed from the knockout drum 1012 and not introduced into the one
or more turboexpanders 1014 can be cooled using one or more heat
exchangers 1018. As noted previously, the heat exchangers may
include any type of heat exchanger suitable for the operation. The
temperature of the portion 1062 of the OCM product gas 1056 can be
decreased using one or more refrigerants, one or more relatively
cool process flows, or combinations of these. The cooled portion
1064 of the OCM product gas 1056 containing a mixture of nitrogen
and methane is introduced into the second separator 1008.
The second separator 1008 may include any system, device or
combination of systems and devices suitable for separating methane
from nitrogen. For example, cryogenic distillation at a relatively
low temperature can be used to promote the separation of liquid
methane from gaseous nitrogen within the second separator 1008. An
exemplary second separator 1008 may include another vertical
distillation column operating significantly below ambient
temperature and above ambient pressure, and also generally below
the temperature of a cryogenic distillation column operating as the
first separator, e.g., as described above. For example, the second
separator 1008 may have an overhead operating temperature of from
about -340.degree. F. (-210.degree. C.) to about -240.degree. F.
(-151.degree. C.); from about -330.degree. F. (-201.degree. C.) to
about -250.degree. F. (-157.degree. C.); about -320.degree. F.
(-196.degree. C.) to about -260.degree. F. (-162.degree. C.); about
-310.degree. F. (-190.degree. C.) to about -270.degree. F.
(-168.degree. C.); or about 300.degree. F. (-184.degree. C.) to
about -280.degree. F. (-173.degree. C.); and a bottom operating
temperature of from about -280.degree. F. (-173.degree. C.) to
about -170.degree. F. (112.degree. C.); about -270.degree. F.
(-168.degree. C.) to about -180.degree. F. (-118.degree. C.); about
-260.degree. F. (-162.degree. C.) to about -190.degree. F.
(-123.degree. C.); about -250.degree. F. (-159.degree. C.) to about
-200.degree. F. (-129.degree. C.); or about -240.degree. F.
(-151.degree. C. to about -210.degree. F. (-134.degree. C.). In
exemplary embodiments, the second separator 1008 will typically
operate at pressures of from about 85 psig (585 kPa) or less; about
70 psig (480 kPa) or less; about 55 psig (380 kPa) or less; or
about 40 psig (275 kPa) or less.
The temperature of at least a portion of the methane-rich effluent
1066 from the second separator 1008 can be increased using one or
more heat exchangers 1020, as described above. After exiting the
one or more heat exchangers 1020, in exemplary embodiments the
temperature of the methane-rich effluent 1066 may be about
125.degree. F. (52.degree. C.) or less; about 100.degree. F.
(38.degree. C.) or less; or about 90.degree. F. (32.degree. C.) or
less, while the pressure of the effluent 1066 may be about 150 psig
(1035 kPa) or less; about 100 psig (690 kPa) or less, or about 50
psig (345 kPa) or less. In an embodiment, e.g., schematically
illustrated in FIG. 10, at least a portion of the methane-rich
effluent 1066 may be recycled back into the feedstock gas 1068 for
the OCM reactor(s) 1002, the feedstock gas/oxygen mixture 1070 the
compressed oxygen containing gas 1072 (from compressor 1028) or
directly to the one or more OCM reactors 1002.
The temperature of at least a portion of the nitrogen-rich effluent
1068 from second separator 1008 can be increased using one or more
heat exchangers 1024 like those described above, such that the
temperature may be raised to about 125.degree. F. (52.degree. C.)
or less; 100.degree. F. (38.degree. C.) or less; or about
90.degree. F. (32.degree. C.) or less, with a pressure of about 150
psig (1035 kPa) or less; about 100 psig (690 kPa) or less; or about
50 psig (345 kPa) or less.
As will be appreciated, in integrating overall systems, while the
one or more heat exchangers 1010, 1016, 1018, 1020, 1022 and 1024
are illustrated as separate heat exchange devices, such heat
exchangers may be integrated into one or more integrated systems,
where the different temperature process flows may be provided in
thermal contact, e.g., as heat exchange media for each other, with
in the heat exchange device or system. In particular, a cooled
process flow that is desired to be heated may be passed through an
opposing portion of a heat exchanger from a heated process flow
that is desired to be cooled, such that the heat from the heated
flow heats the cooler flow, and is, as a result, itself cooled.
Ethylene products of these processes, e.g., in C.sub.2-rich
effluent 1052, are then subjected to additional processing to yield
the desired higher hydrocarbon compositions. For ease of
discussion, the processes and systems for converting ethylene into
higher hydrocarbons are referred to generally as ethylene
conversion processes and systems. A number of exemplary processes
for ethylene conversion are described in greater detail herein.
ETL Integration with Hydrocarbon Processes
The conversion of methane to ethylene, as well as the conversion of
ethylene to higher hydrocarbon compositions, can be carried out in
integrated processes. In some cases, conversion of ethylene to
higher hydrocarbons is performed without conversion of methane to
ethylene. As used herein, integrated processes refers to two or
more processes or systems that are fluidly integrated or coupled
together. Thus, the process for conversion of methane to ethylene
can be fluidly connected to one or more processes for ethylene
conversion to one or more higher hydrocarbon compounds. Fluid
integration or fluid coupling generally refers to a persistent
fluid connection or fluid coupling between two systems within an
overall system or facility. Such persistent fluid communication
typically refers to an interconnected pipeline network coupling one
system to another. Such interconnected pipelines may also include
additional elements between two systems, such as control elements,
e.g., heat exchangers, pumps, valves, compressors, turbo-expanders,
sensors, as well as other fluid or gas transport and/or storage
systems, e.g., piping, manifolds, storage vessels, and the like,
but are generally entirely closed systems, as distinguished from
two systems where materials are conveyed from one to another
through any non-integrated component, e.g., railcar or truck
transport, or systems that are not co-located in the same facility
or immediately adjacent facilities. As used herein, fluid
connection and/or fluid coupling includes complete fluid coupling,
e.g., where all effluent from a given point such as an outlet of a
reactor, is directed to the inlet of another unit with which the
reactor is fluidly connected. Also included within such fluid
connections or couplings are partial connections, e.g., where only
a portion of the effluent from a given first unit is routed to a
fluidly connected second unit. Further, although stated in terms of
fluid connections, it will be appreciated that such connections
include connections for conveying either or both of liquids and/or
gas.
In some instances, a methane to ethylene conversion process is not
just integrated with a single ethylene conversion process, but
instead, is integrated with multiple (i.e., two or more) different
ethylene conversion processes or systems. In particular, ethylene
produced from a single methane feed stream may be converted to
multiple different products using multiple different ethylene
conversion processes. For example, in some embodiments a single OCM
reactor system is fluidly connected to one, two, three, four, five
or more different catalytic or other reactor systems for further
conversion of the ethylene containing product of the OCM reactor
system (also referred to herein as the "ethylene product") to
multiple different higher hydrocarbon compositions.
In some aspects, the ethylene product is selectively directed in
whole or in part to any one or more of the various ethylene
conversion processes or systems integrated with the OCM reactor
system. For example, at any given time, all of the ethylene product
produced through an OCM reactor system may be routed through a
single process. Alternatively, a portion of the ethylene product
may be routed through a first ethylene conversion process or
system, while some or all of the remaining ethylene product is
routed through one, two, three, four or more different ethylene
conversion systems.
Although described in terms of directing ethylene streams to a
single or multiple different ethylene conversion processes, in some
aspects, those ethylene streams may be relatively dilute ethylene
streams, e.g., that contain other components in addition to
ethylene, such as other products of the OCM reaction, unreacted
feed gases, or other by products. Typically, such other components
may include additional reaction products, unreacted feedgases, or
other reactor effluents from an ethylene production process, e.g.,
OCM, such as methane, ethane, propane, propylene, CO, CO.sub.2,
O.sub.2, N.sub.2, H.sub.2, and/or water. The use of dilute ethylene
streams, and particularly those containing other hydrocarbon
components can be particularly advantageous in the ethylene
conversion processes. In particular, because these ethylene
conversion processes can utilize more dilute and less pure streams,
the incoming ethylene streams may not be required to go through as
stringent a separations process or processes as may be required for
other processes intended to produce higher purity ethylene, e.g.,
cryogenic separations systems, lean oil separators, TSA and PSA
based separations processes. These separations processes typically
have relatively high capital costs that scale, at least in part,
based upon the volume of incoming gases. As such, separation
processes for highly dilute ethylene streams can have substantially
high capital and operating costs associated with them. By providing
less stringent separations requirements on these ethylene streams,
one can substantially reduce the capital costs. Further, because
the ethylene conversion processes used in conjunction with the
invention typically result in the production of desired liquid
hydrocarbons, subsequent separation of gas co-products, or
unreacted feed gases is made much simpler.
In addition to reducing capital and operating costs, the use of
ethylene streams that comprise additional hydrocarbon components
can enhance the product slate emanating from the ethylene
conversion processes through which those ethylene streams are
routed. In particular, the presence of higher order hydrocarbons,
C.sub.3, C.sub.4, C.sub.5, etc. in the ethylene streams entering
into the ethylene conversion processes can improve the overall
efficiency of those processes, by providing enriched starting
materials, and also affects the overall carbon efficiency of the
OCM and ethylene conversion processes, by ensuring that a greater
fraction of the carbon input is converted to higher hydrocarbon
products.
While ethylene streams being routed to the ethylene conversion
processes of the invention may range anywhere from trace
concentrations of ethylene to pure or substantially pure ethylene,
e.g., approaching 100% ethylene, the dilute ethylene streams
described herein may generally be characterized as having anywhere
from about 1% to about 50% ethylene, preferably, between about 5%
and about 25% ethylene, and in further preferred aspects, between
about 10% and about 25% ethylene, in addition to other components.
In other embodiments, the ethylene feed gas comprises less than
about 5% ethylene, for example less than about 4%, less than about
3%, less than about 2% or even less than about 1% ethylene. In some
embodiments, the dilute ethylene product gases employed in the
ethylene conversion processes further comprise one or more gases
which are either produced during the OCM reaction or are unreacted
during the OCM process. For example, in some embodiments the
product gas comprises ethylene at any of the foregoing
concentrations and one or more gas selected from CO.sub.2, CO,
H.sub.z, H.sub.2O, C.sub.2H.sub.6, CH.sub.4 and C.sub.3+
hydrocarbons. In some embodiments, such dilute ethylene feed
gasses, which optionally include one or more of the foregoing gases
are advantageous for use in reactions comprising conversion of
ethylene to higher olefins and/or saturated hydrocarbons, for
example conversion of ethylene to liquid fuels such as gasoline
diesel or jet fuel at higher efficiencies (e.g., from methane) than
previously attainable.
By utilizing dilute ethylene streams to feed into one or more
ethylene conversion processes, one eliminates the need to separate
or purify the ethylene coming into the process, e.g., as a product
of an OCM reaction process. The elimination of additional costly
process steps is particularly useful where the ethylene conversion
processes are used to produce lower margin products, such as
gasoline, diesel or jet fuel or blendstocks for these fuels. In
particular, where the desired product is a lower value product, one
may pass the OCM feed gases directly into one or more ethylene
conversion processes that produce hydrocarbon mixtures that can be
used as gasoline, diesel fuel or jet fuel or their blendstocks.
Such direct passage may be in the absence of any intermediate
purification steps, such as any processes used for the removal of
the above described impurities. Alternatively, it may include
certain purification steps to separate out some or all of the
non-hydrocarbon impurities, e.g., N.sub.2, CO.sub.2, CO, H.sub.2,
etc. The direct passage may avoid any hydrocarbon fractionation,
including removal of any of C.sub.1, C.sub.2, C.sub.3, C.sub.4
compounds, or it may include some fractionation, e.g., to enhance
carbon efficiency. For example, such included fractionation may
include separation of methane and or ethane from the OCM effluent
gas to recycle back to the OCM process. In addition to the
foregoing, the presence of additional components such as CO.sub.2,
H.sub.2O and H.sub.2 in the feed streams may also be expected to
improve catalyst lifetime in the ethylene conversion processes by
reducing deactivation, thereby requiring fewer catalyst
regeneration cycles.
In contrast, where one desires to produce more selectively pure
compounds, e.g., aromatic compounds, one may need to pretreat the
feed gases to remove many of the non-ethylene impurities.
Other components of these dilute ethylene streams may include
co-products of the ethylene production processes, e.g., OCM
reactions, such as other C.sub.2+ hydrocarbons, like ethane,
propane, propylene, butane, pentane, and larger hydrocarbons, as
well as other products such as CO, CO.sub.2, H.sub.2, H.sub.2O,
N.sub.2, and the like.
A variety of different ethylene conversion processes may be
employed in the various aspects of the present invention to produce
higher hydrocarbon materials for use in, e.g., chemical
manufacturing, polymer production, fuel production, as well as a
variety of other products. In particular, the ethylene produced
using the OCM processes may be oligomerized and/or reacted by a
variety of different processes and reactor systems for producing
linear alpha-olefins (LAOs), olefinic linear and/or olefinic
branched hydrocarbons, saturated linear and/or branched
hydrocarbons, saturated and/or olefinic cyclic hydrocarbons,
aromatic hydrocarbons, oxygenated hydrocarbons, halogenated
hydrocarbons, alkylated aromatics, and/or hydrocarbon polymers.
In some cases, an ETL sub-system configured to perform an ethylene
conversion process (e.g., oligomerization) can be located between
two OCM sub-systems. A first OCM sub-system produces a first OCM
effluent comprising ethylene and other olefins (e.g., propylene).
This first OCM effluent can be fed into the ETL sub-system, wherein
olefins (e.g., ethylene, propylene) are oligomerized and converted
into higher hydrocarbon products. These higher hydrocarbon
oligomerization products can be recovered from the ETL effluent.
The ETL effluent can be fed into a second OCM sub-system, where
unreacted methane can be converted into ethylene and other olefins
(e.g., propylene) in a second OCM process. The reduced content of
C.sub.2 and C.sub.3 compounds in the second OCM feed stream (due to
the consumption of C.sub.2 and C.sub.3 compounds in the ETL or
oligomerization process) can result in decreased OCM side reactions
and increased C.sub.2 selectivity in the second OCM sub-system. The
effluent from the second OCM sub-system can be processed in a
variety of ways, including by separation in a separations system
(e.g., a 3-way cryogenic separations system). A separations system
can be used to recover methane, which can be recycled to the first
or second OCM system, olefin products (e.g., ethylene, propylene),
and gases such as N.sub.2, CO, and H.sub.2.
ETL processes can result in products such as C.sub.3 and C.sub.4
products, including olefins. Rather than being flared or used for
fuel value, these light olefins can be oligomerized into
higher-value products. C.sub.3 olefins, C.sub.4 olefins, or a
combination thereof can be recovered from an ETL product stream as
a light olefin fraction. This light olefin fraction can be reacted
in a separate oligomerization reactor to produce higher molecular
weight olefins, such as those in the C.sub.6 to C.sub.16 range. The
molecular weight range of the oligomerization products can be tuned
by appropriate choice of catalyst and process parameters. This
approach provides the benefit of increased yield of higher
molecular weight products, including products that are non-aromatic
and mostly olefinic. This oligomerization process can result in
little or no coke formation or other deactivation mechanisms. This
oligomerization process can operate at a temperature of at least
about 50.degree. C. This oligomerization process can operate at a
temperature of at most about 200.degree. C. This oligomerization
process can operate at a temperature from about 50.degree. C. to
about 200.degree. C. Oligomerization catalysts useful for this
process include strong Lewis acid catalysts, AlCl.sub.3/water
solutions, solid superacids, and other solid acid catalysts capable
of oligomerizing C3 and C4 olefins. C4 olefins in this process can
also be used with an alkylation process to generate iso-octane.
Reactor configurations useful for this process include slurry bed,
fixed bed, tubular/isothermal, moving bed, and fluidized bed
reactors, including those disclosed herein (e.g., FIG. 4).
OCM processes, ETL processes, and combinations thereof (e.g.,
OCM-ETL) can result in methane (e.g., unreacted methane), ethane,
and C.sub.3+ non-olefinic hydrocarbon compounds. These compounds
(e.g., methane, ethane, propane, and combinations thereof) can be
converted into aromatic hydrocarbons. For example, excess methane
and ethane from an OCM process can be converted into aromatics with
the use of a catalyst appropriate for ETL, such as those discussed
in this disclosure. Such catalysts can be doped with compounds
including but not limited to molybdenum (Mo), gallium (Ga),
tungsten (W), and combinations thereof. These conversions to
aromatic products can occur in an ETL reactor, as described, and
can also involve the conversion of ethylene to aromatic
products.
Integration of ETL and/or OCM-ETL with Hydrocarbon Processes
The present disclosure provides methods for integrating ETL
sub-systems (or modules) with OCM sub-systems. Such integration can
advantageously enable the formation of products that can be
tailored for various uses, such as, for example fuel. Such
integration can enable the conversion of ethylene in a C.sub.2+
product stream from an OCM reactor to be converted to higher
molecular weight hydrocarbons.
OCM, ETL and OCM-ETL methods and systems of the present disclosure
can be used in greenfield and brownfield contexts. For example, in
a brownfield investment initiative, an OCM-ETL system can be
installed in an old oil refinery. As another example, in a
greenfield investment initiative, an OCM-ETL system can be
installed in a new parcel of round that has access to natural gas.
Brownfield and greenfield initiatives of OCM can be used to meet
world scale production of ethylene.
The present disclosure can be used to form ethylene for various
uses, such as liquefied natural gas (LNG) integration. Liquefied
natural gas (LNG) can be used to enable simplified transport of
natural gas with its volume reduced by at least about 100.times.,
200.times., 300.times., 400.times., 500.times., or 600.times. as
enabled by cooling from a vapor to a liquid. LNG facilities can
include several main process areas--gas treatment area where
natural gas has acid gases, water, and mercury removed; NGL
extraction area; NGL fractionation area; and LNG liquefaction and
storage areas. C.sub.4+ products from raw natural gas are typically
recovered in the fractionation area.
Substantial capital reduction and improved mixed C.sub.4 and
C.sub.5+ yields can be realized by integrating an OCM/ETL process
train into the traditional LNG facility. An OCM-ETL process may be
added to LNG facility such that it utilizes a portion of the main
gas stream that has passed through gas treatment and NGL
extraction. Additional feedstock to the post-bed reactor section of
the OCM may include ethane and propane both fed from high purity
streams generated in the NGL fractionation area. The generated
additional mixed C.sub.4 and C.sub.5+ products can be recovered
though using available capacity in the NGL fractionation process
area. The C.sub.4 and C.sub.5+ products can be referred to as
"SBOB," which may be any composition similar to RBOB but not
meeting one or more ASTM standards.
Alternatively, the process may utilizes the above gas streams (main
gas stream, ethane and propane from fractionation area) in addition
to a dilute methane containing stream produced by the nitrogen
rejection unit in the LNG liquification area. LNG plants can
utilize this stream as low BTU fuel gas for internal energy
generation. The product gas from the OCM-ETL process area can be
fed back into precool portion of the NGL extraction thus enabling
proper extraction of the C.sub.3+ components as well as generation
of a similar low BTU fuel gas. The mixed C.sub.4 and C.sub.5+
(SBOB) products can be recovered though using available capacity in
the NGL fraction process area.
In an example, an OCM-ETL process area can include OCM reaction and
heat recovery areas, process gas compression, ETL guard bed and
reaction section. Steam generated in the OCM heat recovery area can
be effectively used in a refrigerant compressor area to reduce
external power usage.
An OCM-ETL system can be integrated with an LNG facility. To meet
the LNG specification any heavy hydrocarbon (HHC) (e.g., butanes,
pentanes and higher molecular weight) in the natural gas stream may
be removed. However, with the natural gas being dry, it may be
difficult to remove the HHC. Hence an LNG plant may require some
HHC removal process. In some examples, liquids from an ETL system
can be extracted in the pre-cooling process section of the LNG
plant, requiring little to no additional investment. This may
eliminate any additional NGL recovery system that may be required
in a standalone OCM/ETL facility. In addition, a large amount of
steam can be produced in the OCM/ETL system which can be
effectively utilized as shaft power for the refrigerant
compressors. FIG. 11 shows an example of NGL extraction in an LNG
facility. The front-end NGL extraction unit 1100 can take natural
gas 1105 and separate out the natural gas liquids (NGLs) 1110. The
remaining methane 1115 can be compressed 1120 and cooled 1125 to
provide liquified natural gas (LNG) 1130.
FIG. 12 shows an integrated OCM-ETL system for use in LNG
production. The system includes a gas treatment unit, a downstream
NGL extraction unit, a liquefaction unit, and an OCM/ETL sub-system
that generates olefins, such as ethylene. The direction of fluid
flow is indicated by the arrows. The system of FIG. 12 can provide
increased C.sub.5+ and mixed C.sub.4 production. The system of FIG.
12 is exemplary of a typical gas processing plant 1200, with an
OCM-ETL integration. The system contains a gas treatment unit 1202
which takes the incoming natural gas feed stream 1201. The gas
treatment unit can comprise one or more of an acid gas removal
unit, a dehydration unit, mercury removal unit, sulfur removal
unit, or other treatment units. In some cases the acid gas removal
unit is an amine unit, a pressure swing adsorption (PSA) unit, or
another CO.sub.2 removal system. In some cases, the dehydration
unit can be a glycol based water removal unit, a pressure swing
adsorption (PSA) unit and may include a series of separators. A
mercury removal unit can comprise a molecular sieve or an activated
carbon based system. The gas treatment unit can also comprise a
nitrogen removal unit (NRU). An NRU can employ a cryogenic process
or an absorption or an adsorption based process. The treated
natural gas feed 1203 is fed to the NGL extraction unit 1204, which
separates the NGL stream 1205 and a condensate stream containing
heavier hydrocarbons 1210. The heavier hydrocarbons may be C.sub.4+
hydrocarbons. The NGL extraction unit can comprise an adsorption
process unit, a cooling unit (cooling achieved, for example, either
by Joule Thompson expansion, methanol or glycol refrigeration, or a
turboexpander), or a lean oil absorption unit. A part of the stream
1205 is fed to the OCM-ETL reactor system 1207 as stream 1208,
where the methane contained in the stream 1208 is converted to
heavier liquid hydrocarbons via OCM and subsequent ETL conversion
in the reactor system 1207. The liquid hydrocarbons are fed along
stream 1209 back to the NGL extraction unit to recover the
unconverted methane, and separate the heavier condensates and route
them along stream 1210 to NGL ETL fractionation unit 1211. The NGL
ETL fractionation unit can comprise a series of fractionation tower
units, including but not limited to a depropanizer and a
debutanizer to generate a mixed C.sub.4 product 1214, a C.sub.5+
product 1215. The liquefaction unit 1206 produces LNG product 1213.
The advantage of integrating an OCM-ETL reactor system with an
existing natural gas processing plant is envisaged in this case to
generate much more valuable mixed C.sub.4 and C.sub.5+ products. In
addition, the C.sub.2 and C.sub.3 lighter hydrocarbons from the
fractionation unit 1211 can be recycled to the post bed cracking
(PBC) section of the OCM reactor.
The system of FIG. 12 can be modified for use with a diluted
C.sub.1 (methane) fuel gas stream, as shown in FIG. 13. The system
1300 of FIG. 13 includes a pre-cooling system 1320 upstream of the
NGL extraction unit to extract C.sub.3+ compounds, as well as a
nitrogen rejection unit 1316 downstream of the liquefaction unit.
In addition to the system of FIG. 12, the system in FIG. 13
recycles a stream 1319 to the OCM-ETL reactor system to utilize
more of the methane contained in the natural gas feed. The stream
1319 can comprise a high level of methane and inerts such as
nitrogen. A fuel gas stream 1321 is purged from the precooling
section to avoid the buildup of inerts in the system. The nitrogen
rejection unit can comprise a cryogenic based, absorption based or
adsorption based system.
FIG. 14 shows an example OCM-ETL system comprising OCM and ETL
sub-systems, and a separations sub-system downstream of the ETL
sub-system, where 1418 is a condensed water knockout, 1421 is a
process gas compressor, 1423 is a guard bed to remove impurities
such as acetylene and butadiene, 1415 and 1430 are heat recovery,
1437 is a secondary gas compressor, and 1439 is a low temperature
separator. The system in FIG. 14 takes in a treated natural gas
feed stream 1410 and oxygen 1411 from an air separation unit (ASU)
or pipeline, and reacts them in the OCM reactor 1413 to generate an
olefin rich stream 1414 which is then sent to an ETL reactor 1425
to be converted to higher hydrocarbons. The system shows the
various subsystems as the compressors, heat recovery systems
utilizing the high heat of reaction to generate steam and run the
compressors on the steam produced. The system can comprise a
methanation reactor 1427 to convert any CO and CO.sub.2 produced
back to methane, hence adding to the methane content of the sales
gas. The separation subsystem 1403 can comprise a low temperature
separator 1439 to generate lighter methane rich components. The
debutanizer 1442 separates the heavy hydrocarbon condensate to a
C.sub.4 and C.sub.5+ product.
FIG. 15 shows an OCM-ETL system comprising OCM 1906 and ETL 1907
sub-systems, and a cryogenic cold box 1504 downstream of the ETL
sub-system. The OCM-ETL system includes various separations units
for separating C.sub.3 1521 and C.sub.4 1524 components from an ETL
product out of the ETL sub-system. The depropanizer 1521 generates
a C.sub.2-rich stream which is recycled back to the sales gas
export 1509. The C.sub.2 recycle may also be added to the OCM-ETL
reactor subsystem via a recycle stream 1522. The debutanizer
produces a C.sub.4 product 1525 and a C.sub.5+ product 1526.
Refrigeration for the cryogenic cold box can be provided by natural
gas expansion 1502 from at least about 500 PSI, 600 PSI, 700 PSI,
800 PSI, 900 PSI, 1000 PSI, 1500 PSI, or 2000 PSI. The system can
also have a methanation reactor (not shown) to further increase the
methane concentration of the sales gas product. FIG. 15 indicates
an approach to thermally integrate the different streams in the
unit. The system can have an external refrigeration system to
provide for the cryogenic requirements of the unit.
FIG. 16 shows another OCM-ETL in an alternative configuration to
that shown in FIG. 14. The system allows for the recycle of various
streams to the OCM reactor to improve the overall conversion. The
methane rich stream 1648 from the low temperature separation unit
1437 and a C.sub.2 rich stream 1412 from the deethanizer 1619 are
recycled to the sales gas compressor 1617 and the OCM reactor 1413
respectively. The incoming natural gas feed 1610 is treated (to
remove one or more of sulfur, mercury, water, or other components)
in a treatment unit 1611 and then sent to the cryogenic unit 1613
to separate the heavier NGL liquids 1614 which are fed to the
deethanizer 1619. The deethanizer 1619 separates the lighter LNG
product 1620 from heavier C.sub.2+ stream 1412. Feed to OCM is
drawn after the cryogenic unit. As in the systems of FIG. 14 and
FIG. 15, the separations subsystem generates a C.sub.4 and C.sub.5+
product.
It may be noted that the systems of FIG. 14, FIG. 15, and FIG. 16
can be integrated with an existing gas processing plant where one
or more of the existing subsystems can be utilized. The utilization
may arise from the fact that the existing subsystems are no longer
used, or have an additional capacity available to allow for the
integration.
OCM-ETL systems of the present disclosure can be integrated into
and combined into conventional NGL extraction and NGL fractionation
sections of a midstream gas plant. Where NGLs in the gas stream are
declining (or gas is dry), the deployment of OCM-ETL can utilize an
existing facility to produce additional liquid streams. The
implementation of OCM-ETL can allow for the generation of on
specification "pipeline gas." The products from the facility can be
suitable for use (or on specification or "spec") as pipeline gas,
gasoline product, hydrocarbon (HC) stream with high aromatic
content and mixed C.sub.4 product.
An OCM-ETL integrated facility can reduce ethane and propane
product quantities to alleviate pipeline (sales-gas) specification
and liquid handling constraints. Utilities and off-sites facilities
can be effectively utilized. For example steam produced in the
process can offset other shaft power requirements. The capacity of
the OCM-ETL Facility can be varied and can be selected to best fit
specific requirements.
FIGS. 17-18 show examples of OCM-ETL midstream integration. Natural
gas 1701 from upstream can be fed to a gas treatment system 1702
and the treated natural gas can be fed to an NGL extraction unit
1704. In FIG. 17, C.sub.2 and C.sub.3 products 1705 from the NGL
extraction unit 1704 can be directed to an OCM-ETL system 1707 to
generate olefins (e.g., ethylene) and liquids 1709 from the
olefins. Any excess or extracted methane can be directed for use as
pipeline gas 1713. C.sub.4+ hydrocarbons 1710 from the NGL
extraction unit can be directed to an NGL product fractionation
unit for separation in a separations system 1711 into mixed C.sub.4
1714 and C.sub.5+ 1715 product streams, with light hydrocarbons
1712 recycled to the OCM-ETL system 1707. In the system 1800 of
FIG. 18, methane from other natural gas sources 1816 is directed to
a gas conditioning unit 1817 (e.g., to remove sulfur compounds) and
subsequently directed to the OCM-ETL system 1707. Excess methane
1808 can be used as pipeline gas 1713, as an additional feed 1806
into the OCM-ETL system, or both.
Oxygen feed for an OCM unit in the OCM-ETL system can be provided
from air, such as using an air separation unit (e.g., cryogenic air
separation unit), or from an oxygen source, such as pipeline
oxygen.
OCM-ETL systems provided herein can be integrated into a pipeline
NG source, or as part of a new gas processing plant installation
that may provide a source of NG. NG can be provided from a NG
pipeline and/or from a non-OCM process.
In some cases, gas that has been depleted of recoverable
hydrocarbon liquids can be recompressed to a pressure from about
700 PSI to 1500 PSI, or 800 PSI to 1300 PSI, or 900 PSI to 1200 PSI
and returned to the pipeline from which it was initially skimmed.
As an alternative or in addition to, gas that has been depleted of
recoverable hydrocarbon liquids can be recompressed as needed and
piped to downstream of cryogenic gas processing plant. As another
alternative or in addition to, gas that has been depleted of
recoverable hydrocarbon liquids may not be recompressed and is
piped to power plant. As another alternative or in addition to, gas
that has been depleted of recoverable hydrocarbon liquids can be
recompressed as needed and is piped to ammonia plant for use as
synthesis gas blending feedstock. As another alternative or in
addition to, gas that has been depleted of recoverable hydrocarbon
liquids can be recompressed as needed and piped to a methanol plant
for use as synthesis gas feedstock blending.
FIG. 19 shows OCM-ETL systems with various skimmer and recycle
configurations, including a standalone skimmer (top left), a hosted
skimmer (bottom left), a standalone recycle (top right), and a
hosted recycle (bottom right). Under the skimmer configurations,
operation is a once-through process (feed moves forward) where all
feed streams exit the system as product or effluent without
recycle. Under the recycle configurations, some or all of the NG
feed stream is exposed to the OCM catalyst multiple times (feed
moves backward). Such configurations can be employed in stand-alone
settings, in which all or substantially all unit operations are for
OCM/ETL purposes, or hosted settings, in which the unit operations
of an existing non-OCM system at least partially support the
OCM-ETL system. The configurations of FIG. 19 can be used to with a
NG feed of at least about 10 millions of cubic feet per day
(mmcfd), 20 mmcfd, 30 mmcfd, 40 mmcfd, 50 mmcfd, 100 mmcfd, 200
mmcfd, 300 mmcfd, 400 mmcfd, 500 mmcfd, or 1000 mmcfd.
In some cases, depending on economic considerations, various
separations intensities may be utilized for additional liquid
product recovery. In an example, process gas that remains after a
primary liquid recovery section is not further processed and is
returned. In another example, process gas that remains after
primary liquid recovery is fed to a Low Temperature Separator (LTS)
unit where additional hydrocarbon liquids are recovered, such as
C.sub.4+. Effluent gas from this LTS is then returned, such
recycled as described elsewhere herein.
In some cases, process gas that remains after primary liquid
recovery can be fed to a coldbox-based cryogenic unit where
additional hydrocarbon liquids are recovered, such as C.sub.4+.
This coldbox-based cryogenic unit may not utilize deep cryogenic
temperatures and may not require traditional unit operations of
demethanizer and deethanizer. Effluent gas from this unit may then
be returned as described elsewhere herein. In some situations, a
debutanizer column may be installed to provide RVP control of final
C.sub.4+ product and the additional C.sub.4 stream.
There are a number of scenarios where it may be necessary to flare
gas from a midstream gas gathering and or gas processing facility.
In some cases, gas from a midstream system can be burned due to
operating constraints, production gas fluctuations that may result
in capacity limitations of gas gathering and processing facilities,
feed gas conditions that may prevent the gas being processed and
meet the gas specifications, and/or process gas conditions that may
not allow co-processing in a gas facility.
ETL systems of the present disclosure can be integrated in various
existing systems, such as petroleum refineries and/or petrochemical
complexes. Such integration can be with or without OCM systems.
Petroleum refineries and petrochemical complexes may generate a
significant amount and number of purge and other waste gas streams
that may be burned for power generation at fuel value due to their
inability to further process or recover the hydrocarbons. These
waste gas streams may contain a mixture of inert gases and
hydrocarbons, such as olefinic species. An ETL process can be
integrated into a refinery or petrochemical complex such that it
consumes one or several of these olefinic streams and chemically
converts them to higher value oligomers, such as mixed C4 and
C.sub.5+ mixtures. A wide range of these ETL feedstocks can be
generated within a refinery complex.
Several examples of waste gas streams with suitable olefins
include, without limitation, absorber tower overhead product gas
stream containing ethylene and propylene at moderate levels
generated in the light ends recovery process area or the
deethanizer overhead stream containing similar olefins. Streams can
be reacted individually or blended as desired to meet ETL reactor
inlet gas requirements.
An ETL process can include feedstock gas treatment, process gas
compression, ETL reaction and heat recovery sections and units. The
ETL reactor effluent can be returned to the refinery separations
units if existing unit operations and capacity are available. If no
such capacity is available, a small ETL separations sub-system may
be provided to recover the effluent C.sub.4 and C.sub.5+ product
streams using a range of process intensity separations
methods--cooling water or refrigerated condensation recovery,
sponge oil systems, and shallow-grade cryogenic units for the
deepest recovery. Additional process units may be added for further
chemical recovery benefits, including membrane separations to
recovered hydrogen from the ETL reactor effluent post hydrocarbon
recovery.
FIGS. 20-22 show various examples of ETL integration in refineries.
Such systems can employ existing fractionation systems of
refineries to effect product separation.
With reference to FIG. 20, gas from cracking or other units 2001
is, in a refinery gas plant 2002, generated into C.sub.3 and
C.sub.4 products 2004 that are directed to an ETL system 2005,
which generates higher molecular weight hydrocarbons 2006 that are
directed to a product separation system 2007. The product
separation system 2007 can employ an existing separation system of
the refinery. The direction of fluid flow and separations systems
can be selected to effect a given product distribution, such as a
C.sub.2- fuel gas 2008, a C.sub.3 product 2009, and a C.sub.4+
stream 2010. The products 2013 from the fractionation unit 2011 can
be treated to produce a gasoline blend component 2017 and the
heavier products 2012 can be sent to the refinery aromatics
separation unit 2016. Ample integration/blending opportunities
exist in a typical refinery complex.
The systems of FIG. 20 can include a heat exchange ethane cracker
(HXEC) 2119, as shown in FIG. 21. The HXEC can use heat from
flue-gas 2118 to crack ethane to ethylene. The HXEC can utilize one
or more waste heat stream (as during FCC regeneration stage) to
thermally crack the ethane to generate an additional olefin rich
stream as a feed to the ETL reactor. In some cases, the concept can
be used to crack propane feed to produce an olefin rich stream. For
example, removal of coke from catalyst by combustion can generate a
hot flue gas, in some cases with the use of a co-boiler. Flue gas
can reach temperatures of 1600.degree. F. (.about.870.degree. C.),
1800.degree. F. (.about.980.degree. C.), or higher. Heat can be
transferred to a stream comprising ethane, propane, or a
combination thereof, for example in a heat exchanger. This heat can
be used to crack the ethane to ethylene or the propane to
propylene. These olefin products can be used in other processes,
such as in ETL.
A refinery gas plant 2202, receiving gas 2201 from cracking or
other units, can be retrofitted with an OCM-ETL system, as shown in
FIG. 22. The OCM reactor 2212 includes a post-bed cracking (PBC)
unit. The OCM reactor 2212 accepts methane through a natural gas
feed 2211 and generates a product stream 2213 that is directed to a
C.sub.1 separator 2207 that removes methane (C.sub.1) 2208 from the
product stream to provide C.sub.2+ compounds 2214. The methane can
be recycled to the OCM reactor 2210 or directed for use as refinery
fuel 2209. C.sub.2+ compounds 2214 directed to the ETL reactor 2215
are used to generate higher molecular weight hydrocarbons 2216,
which are directed to the refinery gas plant 2202. C.sub.3+
compounds 2204 from the refinery gas plant are directed to a
production fractionation system 2217 for separation. The system of
FIG. 22 can include an HXEC 2223 coupled with the ETL reactor 2215.
The HXEC can ethane 2222 to ethylene 2224, which can then be
directed to the ETL reactor 2215. In some situations, the HXEC is
precluded.
The integration of an ETL system into a refinery or petrochemical
facility can include a cracker and in some cases be performed
without OCM. This can enable polymer grade ethylene to be turned
into gasoline, for example.
It is to be noted with respect to the disclosures above pertaining
to systems described in FIG. 17, FIG. 18, FIG. 19, FIG. 20, FIG.
21, and FIG. 22 that the descriptions are indicative and not
limited to the concepts and configurations represented. One or more
of the following can be additionally integrated into systems such
as those described: a methanation reactor, an ethane skimmer, and
various heat integration configurations and optimizations based on
the refinery configuration, product demand and economics. An ETL
reactor system, including an OCM-ETL reactor system, can be a
versatile system with a wide range of configurations to achieve
economic value from the refinery off gases, waste gases and
additional natural gas feed(s).
Ethylene-to-Liquids (ETL) Integration with Natural Gas
Processing
An aspect of the present disclosure provides olefin-to-liquids
systems and methods. An olefin-to-liquids process can be integrated
in a non-OCM process, such as a natural gas liquids (NGL) system.
The olefin-to-liquids process can be an ethylene-to-liquids (ETL)
process. The ETL process can be part of an OCM system, which can
generate olefins (e.g., ethylene) from methane and an oxidizing
agent (e.g., O.sub.2), as described elsewhere herein. The olefins
can be used as feedstock to one or more ETL reactors for the
conversion of olefins to higher molecular weight hydrocarbons,
which can be in liquid form.
Natural-gas processing is typically a complex industrial process
for cleaning raw natural gas by separating impurities and various
non-methane hydrocarbons and fluids to produce pipeline quality dry
natural gas. Most extracted natural gas can contain, to varying
degrees, low molecular weight hydrocarbon compounds. Examples of
such compounds include methane (CH.sub.4), ethane (C.sub.2H.sub.6),
propane (C.sub.3H.sub.8) and butane (C.sub.4H.sub.10). When brought
to the surface and processed into purified, finished by-products,
all of these are collectively referred to as NGL.
Natural-gas processing plants may purify raw natural gas from (a)
underground gas fields and/or (b) from well heads with associated
gas by removing common contaminates, such as water, carbon dioxide
(CO.sub.2) and hydrogen sulfide (H.sub.z S). Some of the substances
which contaminate natural gas have economic value and are further
processed or sold. A fully operational plant can deliver
pipeline-quality dry natural gas that can be used as fuel by
residential, commercial and industrial consumers.
In some embodiments, existing NGL processing and/or fractionation
systems can be integrated with OCM and ETL processes to produce
various hydrocarbons (which may be liquids), such as alkanes,
alkenes, alkynes, alkoxides, aldehydes, ketones, acids (e.g.,
carboxylic acids), aromatics, paraffins, iso-paraffins, higher
olefins, oligomers or polymers. In some examples, such hydrocarbons
include liquefied petroleum gas (LPG), reformulated gasoline
blendstock for oxygen blending (RBOB), and/or gasoline (e.g.,
natural gasoline or premium gasoline), and/or other hydrocarbon
blendstocks commonly fed to refineries or blending terminals (e.g.,
condensate or diluent). LPG can include propane and butane, and can
be employed as fuel in heating appliances and vehicles.
ETL can be used to generate hydrocarbons for various end uses, such
as gasoline for use in machinery (e.g., automobiles and aircraft).
Products of ETL processes of the present disclosure can be employed
for use as gasoline or jet fuel blend stock, for example. In some
examples, ETL can be used to generate benzene, toluene,
ethylbenzene, and xylenes (BTEX).
An ETL-gasoline process can have a heat of reaction from about 80
KJ/mole to 100 KJ/mol. An adiabatic ETL reactor can have an inlet
temperature of at least about 200.degree. C., 210.degree. C.,
220.degree. C., 230.degree. C., 240.degree. C., 250.degree. C.,
260.degree. C., 270.degree. C., 280.degree. C., 290.degree. C.,
300.degree. C., 310.degree. C., 320.degree. C., 330.degree. C.,
340.degree. C., 350.degree. C., 360.degree. C., 370.degree. C.,
380.degree. C., 390.degree. C., 400.degree. C., or 500.degree. C.
In the reactor, temperature can increase by at least about
50.degree. C. to 150.degree. C., 60.degree. C. to 120.degree. C.,
or 75.degree. C. to 100.degree. C., and process pressure can
increase by at least about 1 bar, 2 bars, 3 bars, 4 bars, 5 bars, 6
bars, 7 bars, 8 bars, 9 bars, 10 bars, 20 bars, 30 bars, 40 bars or
50 bars (absolute).
In some cases, the primary products out of an ETL reactor are
olefins, such as a pentene, hexene or heptene. However, other
secondary products are possible, such as aromatics and paraffins.
In an example, an ETL product has a liquids distribution that is
selective towards C5-C10 hydrocarbons with a substantially low
content of benzene and durene. Such a product may be employed for
use as gasoline. In another example, an ETL product has a liquids
distribution that is selective towards BTEX.
In some examples, a raw natural gas (NG) feed stock can be directed
to an OCM-ETL system to generate C.sub.2+ compounds for use in
generating higher molecular weight hydrocarbons, such as
hydrocarbon products described in the context of FIG. 1. Such
hydrocarbons products can be further processed for various end
uses. For example, the hydrocarbon products can include the
constituents of gasoline, and can be combined with ethanol for use
as automobile fuel.
OCM-ETL systems of the present disclosure can be integrated with
NGL systems to produce NGLs and premium quality gasoline, as well
as NGL associated with natural gas feed or any other hydrocarbon
gas feed stocks. In some examples, NGL processing and/or
fractionation or midstream gas facilities or systems can be
integrated with an OCM reactor system, an ETL reactor system,
separations units, compression units, methanation units and or
other processing units, such as those described in U.S. patent
application Ser. No. 14/099,614, filed on Dec. 6, 2013, which is
entirely incorporated herein by reference.
OCM-ETL systems of the present disclosure can advantageously enable
existing NGL processing systems to be retrofitted for use in
producing various hydrocarbons in an efficient and economical
fashion as compared to other systems presently available. In some
examples, existing NGL processing and/or fractionation plants are
integrated with OCM-ETL systems provided herein, in addition to
other systems that may be required for further processing. OCM-ETL
Integration with NGL processing may include a retrofit of recycle
split vapor (RSV), gas sub-cooled process (GSP) processes or any
gas processing technology. The OCM-ETL plant can include other
systems that may be required for further processing. Existing NGL
systems can be retrofitted with OCM-ETL systems provided herein and
configured, for example, to yield a given product distribution
and/or yield. Such integration may consider the spare or full
capacity of an existing NGL processing plant to reduce the retrofit
capital investments and operating expenses. The OCM-ETL system can
be designed to accommodate any spare capacity of the NGL processing
plant, or to accommodate the NGL plant operating as an OCM/ETL
plant at maximum capacity. The feed (or input) to the OCM-ETL may
be a quantity of fresh natural gas, residue gas, or sales gas
(sales gas and pipeline gas are referring to the same natural gas
which means pipeline specification) that meets the OCM inlet gas
specifications and the spare capacity of the NGL plant. A requisite
amount of an oxidizing agent (e.g., air or oxygen) can be used in
the OCM reactor(s) of the OCM-ETL system.
In some examples, the quantity of the NGL products from OCM-ETL
systems of the present disclosure is at least about 0.3, 0.5, 0.1,
1, 1.5, 2, 3, 4, 5, 6, 7, 8, 9, or 10 gallons per 1000 standard
cubic feet (SCF) of inlet natural gas of inlet gas. In some cases,
the quantity of the NGL products from OCM-ETL systems of the
present disclosure is at least about 0.3, at least about 0.5, at
least about 0.1, at least about 1, at least about 1.5, at least
about 2, at least about 3, at least about 4, at least about 5, at
least about 6, at least about 7, at least about 8, at least about
9, or at least about 10 gallons per 1000 standard cubic feet (SCF)
of inlet natural gas of inlet gas. In some cases, the quantity of
the NGL products from OCM-ETL systems of the present disclosure is
at most about 0.3, at most about 0.5, at most about 0.1, at most
about 1, at most about 1.5, at most about 2, at most about 3, at
most about 4, at most about 5, at most about 6, at most about 7, at
most about 8, at most about 9, or at most about 10 gallons per 1000
standard cubic feet (SCF) of inlet natural gas of inlet gas. In
some cases, the quantity of the NGL products from OCM-ETL systems
of the present disclosure is in the range of 0.8 gallons to 1.5
gallons per 1000 standard cubic feet (SCF) of inlet natural gas of
inlet gas.
FIG. 23A shows an NGL process. The NGL system comprises a raw
natural gas feed stream 2301, an NGL system 2302 and a product
stream 2303. The raw natural gas feed stream 2301 comprises methane
(CH.sub.4) in addition to other chemicals (e.g., H.sub.2O, CO.sub.2
and H.sub.2S). The NGL system 2302 can comprise various processing
equipment for refining the feed stream 2301 to generate the product
stream 2303 comprising one or more hydrocarbon products, such as
methane, ethane, propane and/or butane. Such processing equipment
can include separations units, such as distillation columns. In
some examples, the product stream 2303 comprises methane at a
concentration (or purity) that is higher as compared to the feed
stream 2301. The product stream 2303 can be directed to a natural
gas pipeline for distribution of natural gas to end users.
In FIG. 23B, the NGL process of FIG. 23A has been retrofitted with
an OCM-ETL system of the present disclosure. FIG. 23B shows the
feed stream 2301 directed into NGL system 2304. The NGL system 2304
can include at least a subset or all of the equipment of the NGL
system 2302 described in the context of FIG. 23A. In an example,
the NGL system 2304 is the NGL system 2302. The NGL system 2304
yields the product stream 2303 and an additional product stream
2305. The additional product stream 2305 is directed to an OCM-ETL
system 2306. The OCM-ETL system 2306 generates product stream 2307,
which can include C.sub.2+ compounds. In some examples, the product
stream comprises C.sub.3-C.sub.12 hydrocarbons.
Although FIG. 23B shows an NGL process retrofitted with an OCM-ETL
system, other non-OCM processes may be retrofitted with the OCM-ETL
system. For example, the OCM-ETL system can be integrated in an oil
refinery, and products from crude oil refining may be directed to
the OCM-ETL system for further processing.
FIG. 24 shows a system 2400 comprising an existing gas plant 2401
that has been retrofitted with an OCM-ETL system 2402. The OCM-ETL
system 2402 may be used with ethylene or other olefins. A raw
natural gas (NG) feed 2403 is directed into the existing gas plant
2401, which comprises a treatment unit 2404, NGL extraction unit
2405, compression unit 2406 and fractionation unit 2407. The NGL
extraction unit 2405 can be a demethanizer unit, optionally a
demethanizer unit incorporated with a recycle split vapor (RSV)
retrofit or stand-alone unit. The treatment unit 2404 removes water
and CO.sub.2 from the NG feed 2403 and directs natural gas to the
NGL extraction unit 2405. In some cases, the treatment unit removes
sulfur from the NG feed. The NGL extraction unit 2405 removes
methane, ethane, CO.sub.2 and N.sub.2 from the NG feed 2403, and
directs methane, ethane, CO.sub.2 and N.sub.2 to the compression
unit 2406 along fluid stream 2408. At least a portion of the
methane from the fluid stream 2408 is directed along stream 2409 to
an OCM reactor 2410 of the OCM-ETL system 2402. The compression
unit 2406 compresses methane in the fluid stream 2408 and directs
compressed methane to a natural gas pipeline 2411 for distribution
of methane to end users.
With continued reference to FIG. 24, C.sub.2+ compounds from the
NGL extraction unit 2405 are directed to the fractionation unit
2407, which can be a distillation column. The fractionation unit
2407 splits the C.sub.2+ compounds into streams comprising various
C.sub.2+ compounds, such as a C.sub.2 stream 2412 along with
C.sub.3 2423, C.sub.4 and C.sub.5 streams. The C.sub.2 stream 2412
and/or C.sub.3 stream 2423 can be directed to a post-bed cracking
(PBC) unit 2413 of the OCM-ETL system 2402. In some cases, C.sub.3,
C.sub.4 and/or C.sub.4+ compounds are directed to the PBC unit.
Examples of post-bed cracking is described in U.S. patent
application Ser. No. 14/553,795, filed Nov. 25, 2014, which is
entirely incorporated herein by reference.
In the OCM-ETL system 2402, methane from the stream 2409 and air
2414 are directed to the OCM reactor 2410. The OCM reactor 2410
generates an OCM product stream comprising C.sub.2+ compounds in an
OCM process, as discussed elsewhere herein. C.sub.2+ alkanes (e.g.,
ethane) in the product stream, as well as C.sub.2 alkanes in the
C.sub.2 stream 2412, may be cracked to C.sub.2+ alkenes (e.g.,
ethylene) in the post-bed cracking (PBC) unit 2413 (which can be a
downstream component of the OCM reactor 2410). The product stream
is then directed to a condenser 2415, which removes water from the
product stream. The product stream is then directed to a
compression unit 2416 and subsequently a pressure swing absorption
(PSA) unit 2417. The PSA separates N.sub.2, CO, CO.sub.2, H.sub.2O,
H.sub.2 and some methane from C.sub.2+ compounds in the product
stream, and directs the C.sub.2+ compounds to one or more ETL
reactors 2418 of the OCM-ETL system 2402. The stream comprising
nitrogen 2424 (and in some cases CH.sub.4, CO.sub.2, H.sub.2O,
H.sub.2 and/or CO) can be fed into a fuel gas stream for use in
generating power, burning, use as a thermal oxidizer. The C.sub.2+
compounds directed into the ETL reactor 2418 can include ethane,
ethylene, propane, propylene, along with methane, CO, CO.sub.2,
H.sub.2, N.sub.2 and water. The ETL reactor 2418 generates higher
molecular weight hydrocarbons, such as C.sub.4-C.sub.12 (e.g.,
C.sub.4+ or C.sub.5+) compounds (e.g., butane, butylenes, pentane,
hexane, etc.). A product stream from the ETL reactor 2418 is
directed to another compression unit 2419 and subsequently a
vapor-liquid separator (or knock-out drum) 2420, which separates
liquids (e.g., the C.sub.5+ compounds) from vapors (e.g., methane)
in the product stream and provides a product stream 2421 comprising
the C.sub.5+ compounds. Remaining compounds, including vapors
(e.g., methane), are recycled to the feed stream 2403 along recycle
stream 2422. Methane directed along stream 2422 can be directed to
the OCM reactor 2410 for further C.sub.2+ product generation.
The OCM-ETL 2402 system can include one or more OCM reactor 2410.
For example, the OCM reactor 2410 can be an OCM reactor train
comprising multiple OCM reactors. In addition to, or as an
alternative, the OCM-ETL system 2402 can include one or more ETL
reactor 2418. For example, the ETL reactor 2418 can be multiple ETL
reactors in parallel, with each ETL reactor configured to generate
a given hydrocarbon (see, e.g., FIG. 1). In some cases, C.sub.3
and/or C.sub.4 compounds can be taken from the fractionators and
fed into a further downstream region of a post-bed cracking (PBC)
reactor for olefin production.
The compression units 2406, 2416 and 2419 can each be a multistage
gas compression unit. Each stage of such multistage gas compression
unit can be followed by cooling and liquid hydrocarbon and water
removal.
The OCM-ETL system 2402 can be operated with other oxidizing
agents, such as previously separated O.sub.2 such as O.sub.2 from a
pipeline or as a product from an air separation unit (ASU). In the
alternative configuration of FIG. 25, an O.sub.2 feed stream 414 is
directed to the OCM reactor 2410. The O.sub.2 feed stream 2514 can
be generated, for example, using a cryogenic air separation unit
(not shown), which separates air into individual streams comprising
O.sub.2 and N.sub.2. The system of FIG. 25 further includes a
methanation system 2523 (see below) that converts CO, CO.sub.2 and
H.sub.2 from the vapor-liquid separator to methane, which can be
recycled along stream 2422.
In the figures, the direction of fluid flow between units is
indicated by arrows. Fluid may be directed from one unit to another
with the aid of valves and a fluid flow system. In some examples, a
fluid flow system can include compressors and/or pumps, as well as
a control system for regulating fluid flow, as described elsewhere
herein.
Methanation Systems
Oxidative Coupling of Methane (OCM) is a process that may convert
natural gas (or methane) to ethylene and other longer hydrocarbon
molecules via reaction of methane with oxygen. Given the operating
conditions of OCM, side reactions can include reforming and
combustion, which can lead to the presence of significant amounts
of H.sub.2, CO and CO.sub.2 in the effluent stream. Typical H.sub.2
content in the effluent stream can range between about 5% and about
15%, between about 1% and about 15%, between about 5% and about
10%, or between about 1% and about 5% (molar basis). CO and
CO.sub.2 can each range between about 1% and about 5%, between
about 1% and about 3%, or between about 3% and about 5% (molar
basis). In some cases, the ethylene and all the other longer
hydrocarbon molecules contained in the effluent stream are
separated and purified to yield the final products of the process.
This can leave an effluent stream containing the unconverted
methane, hydrogen, CO and CO.sub.2 and several other compounds,
including low amounts of the product themselves depending on their
recovery rates.
In some cases, this effluent stream needs to be recycled to the OCM
reactor. However, if CO and H.sub.2 are recycled to the OCM reactor
along with methane, they can react with oxygen to produce CO.sub.2
and H.sub.2O, causing various negative consequences to the process
including, but not limited to: (a) an increase of the natural gas
feed consumption (e.g., because a larger portion of it can result
in CO.sub.2 generation instead of product generation); (b) a
decrease of the OCM per-pass methane conversion (e.g., because a
portion of the allowable adiabatic temperature increase can be
exploited by the H.sub.2 and CO combustion reactions instead of the
OCM reactions); and an increase of the oxygen consumption (e.g.,
because some of the oxygen feed can react with CO and H.sub.2
instead of methane).
In some instances, the effluent stream is exported to a natural gas
pipeline (i.e., to be sold as sales gas into the natural gas
infrastructure). Given that specifications can be in place for
natural gas pipelines, the concentrations of CO, CO.sub.2 and
H.sub.2 in the effluent can need to be reduced to meet the pipeline
requirements.
In some embodiments, the effluent stream may also be used as a
feedstock for certain processes that may require lower
concentrations of H.sub.2, CO and CO.sub.2.
Therefore, it can be desirable to reduce the concentration of
H.sub.2, CO and CO.sub.2 in the OCM effluent stream, upstream or
downstream of the separation and recovery of the final products.
This can be accomplished using methanation systems and/or by
separating H.sub.2 and CO from the effluent stream (e.g., using
cryogenic separations or adsorption processes). The disclosure also
includes separating CO.sub.2 from the effluent stream using
CO.sub.2 removal processes, such as chemical or physical absorption
or adsorption or membranes. However, these separation processes can
require significant capital investments and can consume
considerable amounts of energy, in some cases making an OCM-based
process less economically attractive.
Described herein are systems and methods for reducing CO, CO.sub.2
and H.sub.2 concentration in a methane stream. The method comprises
reacting these compounds to form methane in a reaction called
methanation.
CO.sub.2 and/or sulfur-containing compounds (e.g., H.sub.2S) can be
separated via a CO.sub.2 removal unit, such as, for example, an
amine-based system, a caustic system or any other physical or
chemical absorption or adsorption unit. CO and H.sub.2 can be
separated together with the methane in a cryogenic separator. If CO
and H.sub.2 are recycled to an OCM reactor along with methane, they
can react with oxygen (e.g., pure O.sub.2 or O.sub.2 in air) to
produce CO.sub.2 and H.sub.2O, causing various negative
consequences to the process, including, without limitation: (i)
increase in natural gas feed consumption and a decrease in C.sub.2+
product generation; (ii) decrease of the OCM per-pass methane
conversion; and (iii) increase in oxygen consumption. Given the
potential negative effects of the presence of CO and H.sub.2 in a
stream comprising methane, it may be preferable to minimize the
concentration of CO and H.sub.2. In addition, by converting the CO
and H.sub.2 back into methane, the carbon efficiency of the process
can be increased by recycling the methane to the OCM reactor or to
the natural gas pipeline.
An aspect of the present disclosure provides a methanation system
that can be employed to reduce the concentration of CO, CO.sub.2
and H.sub.2 in a given stream, such as an OCM product stream as
well as improve carbon efficiency. This can advantageously minimize
the concentration of CO, CO.sub.2 and H.sub.2 in any stream that
may be ultimately recycled to an OCM reactor. The methanation
system can be employed for use with any system of the present
disclosure, such as the OCM-ETL system 302 described above.
In a methanation system, CO reacts with H.sub.2 to yield methane
via CO+3H.sub.2.fwdarw.CH.sub.4+H.sub.2O. In the methanation
system, CO.sub.2 can react with H.sub.2 to yield methane via
CO.sub.2+4H.sub.2.fwdarw.CH.sub.4+2 H.sub.2O. Such processes are
exothermic and generate heat that may be used as heat input to
other process units, such as heating an OCM reactor of a PBC
reactor, or pre-heating reactants, such as methane and/or an
oxidizing agent (e.g., O.sub.2) prior to an OCM reaction.
In some cases, to limit the heat release per unit of flow of
reactants, methanation can be conducted on streams that contain CO,
CO.sub.2, H.sub.2 and a suitable carrier gas. The carrier gas can
include an inert gas, such as, e.g., N.sub.2, He or Ar, or an
alkane (e.g., methane, ethane, propane and/or butane). The carrier
gas can add thermal heat capacity and significantly reduce the
adiabatic temperature increase resulting from the methanation
reactions.
In some examples, methane and higher carbon alkanes (e.g., ethane,
propane and butane) and nitrogen are employed as carrier gases in a
methanation process. These molecules can be present in an OCM
process, such as in an OCM product stream comprising C.sub.2+
compounds. Downstream separation units, such as a cryogenic
separation unit, can be configured to produce a stream that
contains any (or none) of these compounds in combination with CO
and H.sub.2. This stream can then be directed to the methanation
system.
A methanation system can include one or more methanation reactors
and heat exchangers. CO, CO.sub.2 and H.sub.2 can be added along
various streams to the one or more methanation reactors. A
compressor can be used to increase the CO.sub.2 stream pressure up
to the methanation operating pressure, which can be from about 2
bar (absolute) to 60 bar, or 3 bar to 30 bar. CO.sub.2 can be added
to the inlet of the system in order to create an excess of CO.sub.2
compared to the amount stoichiometrically required to consume all
the available H.sub.2. This is done in order to minimize H.sub.2
recycled to OCM, which may not be preferable.
Given the exothermicity of the methanation reactions, a methanation
system can include various methanation reactors for performing
methanation. In some cases, a methanation reactor is an adiabatic
reactor, such as an adiabatic fixed bed reactor. The adiabatic
reactor can be in one stage or multiple stages, depending, for
example, on the concentration of CO, CO.sub.2 and H.sub.2 in the
feed stream to the methanation system. If multiple stages are used,
inter-stage cooling can be performed by either heat exchangers
(e.g., a stage effluent can be cooled against the feed stream or
any other colder stream available in the plant, such as boiler feed
water) or quenching via cold shots, i.e. the feed stream is divided
into multiple streams, with one stream being directed to the first
stage while each of the other feed streams being mixed with each
stage effluent for cooling purposes. As an alternative, or in
addition to, a methanation reactor can be an isothermal reactor. In
such a case, reaction heat can be removed by the isothermal reactor
by, for example, generating steam, which can enable a higher
concentration of CO, CO.sub.2 and H.sub.2 to be used with the
isothermal reactor. Apart from adiabatic and isothermal reactors,
other types of reactors may be used for methanation.
FIG. 26 shows an example methanation system 2600. The system 2600
comprises a first reactor 2601, second reactor 2602 and a heat
exchanger 2603. The first reactor 2601 and second reactor 2602 can
be adiabatic reactors. During use, a recycle stream 2604 comprising
methane, CO and H.sub.2 (e.g., from a cryogenic separation unit) is
directed to the heat exchanger 2603. In an example, the recycle
stream 2604 comprises between about 65% and 90% (molar basis)
methane, between about 5% and 15% H.sub.2, between 1% and 5% CO,
between about 0% and 0.5% ethylene, and the balance inert gases
(e.g., N.sub.2, Ar and He). The recycle stream 2604 can have a
temperature from about 20.degree. C. to 30.degree. C., and a
pressure from about 2 bar to 60 bar (absolute), or 3 bar to 30 bar.
The recycle stream 2604 can be generated by a separation unit
downstream of an OCM reactor, such as a cryogenic separation
unit.
In the heat exchanger 2603, the temperature of the recycle stream
2604 is increased to about 100.degree. C. to 400.degree. C., or
200.degree. C. to 300.degree. C. The heated recycle stream 2604 is
then directed to the first reactor 2601. In the first reactor 2601,
CO and H.sub.2 in the recycle stream 2604 react to yield methane.
This reaction can progress until all of the H.sub.2 is depleted
and/or a temperature approach to equilibrium of about 0 to
30.degree. C., or 0 to 15.degree. C. is achieved. The methanation
reaction in the first reactor 2601 can result in an adiabatic
temperature increase of about 20.degree. C. to 300.degree. C., or
50.degree. C. to 150.degree. C.
Next, products from the first reactor, including methane and
unreacted CO and/or H.sub.2, can be directed along a first product
stream to the heat exchanger 2603, where they are cooled to a
temperature of about 100.degree. C. to 400.degree. C., or
200.degree. C. to 300.degree. C. In the heat exchanger 2603, heat
from the first product stream 2603 is removed and directed to the
recycle stream 2604, prior to the recycle stream 2604 being
directed to the first reactor 2601.
Next, a portion of the heated first product stream is mixed with a
CO.sub.2 stream 2605 to yield a mixed stream that is directed to
the second reactor 2602. The CO.sub.2 stream 2605 can be generated
by a separation unit downstream of an OCM reactor, such as a
cryogenic separation unit. This can be the same separation unit
that generated the recycle stream 2604. In some cases, the methods
described herein increase carbon efficiency compared to methods
that do not use methanation. For example, the amount of CO and/or
CO.sub.2 can be reduced by at least about 5%, at least about 10%,
at least about 20%, at least about 50%, at least about 75% or at
least about 80%.
In the second reactor 2602, CO and CO.sub.2 react with H.sub.2 to
yield a second product stream 2606 comprising methane. The
reaction(s) in the second reactor 2602 can progress until
substantially all of the H.sub.2 is depleted and/or a temperature
approach to equilibrium of about 0 to 30.degree. C., or 0 to
15.degree. C. is achieved. The proportions of CO, CO.sub.2 and
H.sub.2 in the mixed stream can be selected such that the second
product stream 2606 is substantially depleted in CO and H.sub.2. In
some cases, the second product stream 2606 is fed back into the
natural gas feed of a natural gas to liquids facility.
The first reactor 2601 and the second reactor 2602 can be two
catalytic stages in the same reactor vessel or can be arranged as
two separate vessels. The first reactor 2601 and second reactor
2602 can each include a catalyst, such as a catalyst comprising one
or more of ruthenium, cobalt, nickel and iron. The first reactor
2601 and second reactor 2602 can be fluidized bed or packed bed
reactors. Further, although the system 2600 comprises two reactors
2601 and 2602, the system 2600 can include any number of reactors
in series and/or in parallel, such as at least 1, 2, 3, 4, 5, 6, 7,
8, 9, 10, 20, 30, 40, or 50 reactors.
Although the CO.sub.2 stream 2605 is shown to be directed to the
second reactor 2602 and not the first reactor 2601, in an
alternative configuration, at least a portion or the entire
CO.sub.2 stream 2605 can be directed to the first reactor 2601. The
proportions of CO, CO.sub.2 and H.sub.2 can be selected such that
the methanation product stream is substantially depleted in CO and
H.sub.2.
Methane generated in the system 2600 can be employed for various
uses. In an example, at least a portion of the methane can be
recycled to an OCM reactor (e.g., as part of an OCM-ETL system) to
generate C.sub.2+ compounds, including alkenes (e.g., ethylene). As
an alternative, or in addition to, at least a portion of the
methane can be directed to a non-OCM process, such as a natural gas
stream of a natural gas plant (see, e.g., FIGS. 3 and 4). As an
alternative, or in addition to, at least a portion of the methane
can be directed to end users, such as along a natural gas
pipeline.
The methanation reaction can be practiced over a nickel-based
catalyst, such as those used to produce SNG (Substitute Natural Gas
or Synthetic Natural Gas) from syngas or used to purify streams
containing CO and CO.sub.2 (e.g., to remove CO and CO.sub.2 present
in the make-up feed to an ammonia synthesis unit). Examples of such
catalysts include the KATALCO.TM. series (including models 11-4,
11-4R, 11-4M and 11-4MR) that are include nickel supported on
refractory oxides; the HTC series (including NI 500 RP 1.2) having
nickel supported on alumina; and Type 146 having ruthenium
supported on alumina. Additional methanation catalysts include
models PK-7R and METH-134. The methanation catalyst can be tableted
or an extrudate. The shapes of such catalysts can be, for example,
cylindrical, spherical, or ring structures, or partial shapes
and/or combinations thereof. In some cases, ring structures are
advantageous due to their reduced pressure drop across the reactor
bed relative to cylindrical and spherical commercial forms. In some
cases, the methanation catalyst is a doped or modified version of a
commercially available catalyst.
In some cases, merely applying a methanation catalyst to the OCM
and/or ETL process that has been developed or optimized for another
process (e.g., SNG production or gas purification) can result in
operational problems and/or non-optimal performance, including
carbon formation (or coking) over the methanation catalyst. Coking
can lead to de-activation of the catalyst and, eventually, to loss
of conversion through the methanation reactor, thus making the
methanation process ineffective, severely limiting the performances
of the overall OCM and/or ETL-based process and, possibly,
preventing the final products from achieving the required
specifications.
The selectivity and/or conversion produced by an existing and/or
commercially available methanation catalyst at a given process
condition (e.g., gas-hourly space velocity, molar composition,
temperature, pressure) may not be ideal for OCM and/or ETL
implementations. For example, ammonia plants can have between about
100 ppm and 1% CO with a molar excess of H.sub.2 (e.g., 2, 5, 10,
50, 100-fold or more excess) that drives equilibrium in favor of
complete methanation. Methanation systems in ammonia plants have a
small temperature difference between inlet and outlet of the
adiabatic methanation reactor (e.g., 20 to 30.degree. C.) and can
be sized for catalyst lifetime. SNG production does not have a vast
molar excess of H.sub.2 in some cases. Methanation in SNG processes
can have an inlet versus outlet temperature difference of greater
than 100.degree. C. and be performed in multiple stages.
Furthermore, the purpose of methanation can be different for OCM
and/or ETL. Ammonia and SNG processes typically perform methanation
primarily to eliminate CO and/or CO.sub.2 (although H.sub.2 can
also be eliminated or substantially reduced in concentration),
while methanation is performed in OCM and/or ETL processes
primarily to eliminate H.sub.2 (although CO and/or CO.sub.2 can
also be eliminated or substantially reduced in concentration).
A methanation catalyst and/or catalytic process is described herein
that can prevent or reduce carbon formation in the methanation
reactor or other operational inefficiencies. The catalyst and/or
catalytic process can be achieved through any combination of: (a)
removing chemical species that can contribute to coke formation
from the methanation inlet feed; (b) introducing chemical species
into the methanation feed that eliminate or reduce the rate of coke
formation; and (c) using the methanation catalyst described herein
that reduces or eliminates coke formation and/or is designed to
operate at the process conditions of OCM and/or ETL effluent or OCM
and/or ETL process streams (e.g., gas-hourly space velocity, molar
composition, temperature, pressure).
In some instances, the species present in the OCM and/or ETL
effluent stream that can lead to carbon formation in the
methanation reactor are removed or reduced in concentration using a
separations or reactive process. The typical operating conditions
of a methanation reactor can be between about 3 and about 50 bar
pressure and between about 150 and about 400.degree. C.
temperature. Any hydrocarbon species containing carbon-carbon
double or triple bonds is sufficiently reactive to form carbon
deposits (i.e., coke). Examples of these species are acetylene, all
olefins and aromatic compounds. Removal or significant reduction of
these species can be achieved via different methods including, but
not limited to: (a) hydrogenation (i.e., reaction of these species
with the hydrogen present in the effluent stream itself to produce
alkanes) over suitable catalysts prior to the methanation reactor;
(b) condensation and separation of these species from methane prior
to the methanation reactor; (c) absorption or adsorption of these
species; (d) by utilizing suitable membranes; or (d) any
combination thereof.
In embodiments of the present disclosure, new species are
introduced into the methanation inlet stream that eliminate or
reduce the rate of carbon formation. Molecular species that can
create a reducing atmosphere can be used to counteract an oxidation
reaction and can therefore reduce the rate of carbon formation.
Hydrogen and water are examples of these particular compounds and
can be added to the OCM and/or ETL effluent stream prior to
methanation to increase their concentration in the methanation
reactor.
An aspect of the present disclosure provides a methanation catalyst
for an OCM and/or ETL process. Coke formation is typically the
product of surface driven reactions. Therefore, the methanation
catalyst for OCM and/or ETL alters the local electronic environment
around the active site of the catalyst. This can be done by
changing the elemental composition (for example via
post-impregnation doping, or creating a new mixed metal of nickel
and another transition metal), morphology and structure (for
example via synthesizing the metal in a non-bulk form factor).
Examples of such syntheses include; nanowires of the same material,
nanoparticles coated on a support, and vapor deposition of the
active material on a support material. Additional modifications to
the surface may result from post synthetic processing steps, such
as etching of the surface, oxidizing and reducing the metal to
create a different surface reconstruction, calcination steps under
different atmospheres (e.g., oxidizing or reducing), heating to
achieve different crystal phases, and inducing defect formation.
The end result of said modifications of the methanation catalyst is
specifically designed to minimize carbon (coke) formation, while
still effectively at conducting the methanation reactions.
The methanation process and/or methanation catalyst operates with
OCM and/or ETL product gas, either directly or after one or more
heat exchangers or separation operations. For example, the
methanation feed stream can have the following composition on a
molar basis: CH.sub.4 between about 65% and about 90%; H.sub.2
between about 5% and about 15%; CO between about 1% and about 5%
(molar basis); C.sub.2H.sub.4 between about 0% and about 0.5%; and
C.sub.2H.sub.2 between about 0% and about 0.1%. As described
herein, the ETL effluent can contain C.sub.2+ compounds including
propane, propylene, butane, butylene, and C.sub.5+ compounds. These
C.sub.2+ compounds can be present in the stream entering the
methanation reactor in any concentration. The balance of the feed
stream can be inert gases such as N.sub.2, Ar and He. The
methanation feed stream typically has a temperature close to
ambient temperature and a pressure ranging between about 3 and
about 50 bar.
In some cases, the entire ETL product stream and/or all of the
C.sub.2+ compounds present in the ETL effluent and/or any or all of
the olefins present in the ETL effluent are fed into the
methanation reactor (i.e., methanation feed). In some cases, the
temperature of the ETL effluent is not reduced, or not
substantially reduced before being fed into the methanation reactor
such that all or most (at least about 70%, at least about 80%, at
least about 90%, at least about 95%, or at least about 99%) of the
C.sub.2+ compounds and/or olefins remain in the methanation feed.
In some cases, the temperature of the ETL effluent is reduced to
separate some of the C.sub.2+ compounds and/or olefins from the
stream before being fed into the methanator. The temperature can be
reduced to a temperature sufficiently low to remove most (at least
about 70%, at least about 80%, at least about 90%, at least about
95%, or at least about 99%) of the C.sub.5+ compounds (e.g., about
40.degree. C.), to remove most (at least about 70%, at least about
80%, at least about 90%, at least about 95%, or at least about 99%)
of the C.sub.4+ compounds (e.g., about 10.degree. C.), or to remove
most (at least about 70%, at least about 80%, at least about 90%,
at least about 95%, or at least about 99%) of the C.sub.3+
compounds (e.g., about -40.degree. C.).
The methanation reaction can produce water and/or have water in the
methanation effluent. In some cases, it can be desirable to remove
this water prior to recycling the methanation effluent to the OCM
reactor. This can be accomplished by lowering the temperature of
the methanation effluent or performing any separation procedure
that removes the water. In some embodiments, at least about 70%, at
least about 80%, at least about 70%, at least about 90%, at least
about 95%, or at least about 99% of the water is removed from the
methanation effluent prior to the OCM reactor. Removing the water
can increase the lifetime and/or performance of the OCM
catalyst.
In some cases, the ETL process can be designed to use a methanation
catalyst that is not optimized for the ETL process. The OCM or
OCM-ETL process can be designed to produce gasoline or distillates
or aromatics (or any combination thereof) from natural gas. In this
case, the effluent of the OCM reactor is fed to an ETL reactor
where all short olefins (e.g., ethylene and propylene) are
converted to longer chain hydrocarbons over a suitable
oligomerization catalyst. An example of such a catalyst is the
zeolite ZSM-5. The product stream that contains unconverted
methane, unconverted olefins, CO, CO.sub.2, H.sub.2, water, inert
species and all oligomerization products (paraffins, isoparaffins,
olefins and aromatics) is fed to the methanation module. The
concentration of the oligomerization products in the methanation
feed stream can vary depending on the type and extent of separation
conducted prior to the methanation step. The methanation feed
stream typically has a temperature close to or below ambient
temperature and a pressure ranging between 3 and 50 bar.
With reference to FIG. 27, the methanation system can be designed
to use a catalyst that is not necessarily optimized for the OCM
and/or ETL process streams. The methanation feed stream 2700 is
first sent to a first heat exchanger 2705 where its temperature is
increased to the methanation reactor inlet temperature, typically
between 150 and 300.degree. C. Steam is injected 2710 immediately
downstream of the heat exchanger to increase water concentration in
the methanation feed stream. Then the heated stream is fed to a
first adiabatic reactor 2715 where ethylene, acetylene and any
other hydrocarbon that presents multiple carbon-carbon bonds are
hydrogenated via reaction with the H.sub.2 present in the stream
itself.
The effluent from 2715 is then fed to a second reactor 2720, where
CO and CO.sub.2 react with H.sub.2 until a desired approach to
equilibrium is achieved, typically 0-15.degree. C. to equilibrium.
The adiabatic temperature increase that results from CO and
CO.sub.2 methanation can depend on the composition of the feed
stream, and is typically in the 50-150.degree. C. range.
The effluent from the second reactor 2720 is then sent to the first
heat exchanger 2705 and a second heat exchanger 2725 where it is
cooled down to a temperature below water condensation. The stream
is then fed to a phase separator 2730 where the condensed water and
a portion of the longer hydrocarbons is separated from the
vapors.
The vapor stream from the phase separator 2735 is sent to the final
product purification and recovery section or injected into a
natural gas pipeline, depending on its concentration.
Alternatively, the vapor stream 2735 from the phase separator 2730
can be further methanated in a second methanation reactor to
further reduce CO, CO.sub.2 and H.sub.2 concentration (not
shown).
The liquid stream from the phase separator 2740 is re-injected into
the methanation feed stream alongside the steam. Alternatively, it
can be first vaporized and then re-injected, or it can be sent to a
water treatment system for water recovery and purification (not
shown).
The reactors 2715, 2720 (and a third reactor, if present) or any
combination of them can be physically situated in the same vessel
or can be arranged in separate individual vessels.
In processes, systems, and methods of the present disclosure, a
Fischer-Tropsch (F-T) reactor can be used to replace a methanation
reactor, for example in a methane recycle stream. CO and H.sub.2,
such as that found in a methane recycle stream, can be converted to
a variety of paraffinic linear hydrocarbons, including methane, in
an F-T reaction. Higher levels of linear hydrocarbons, such as
ethane, can improve OCM process efficiency and economics. For
example, effluent from an OCM reactor can be directed through a
cooling/compression system and other processes before removal of a
recycle stream in a de-methanizer. The recycle stream can comprise
CH.sub.4, CO, and H.sub.z, and can be directed into an F-T reactor.
The F-T reactor can produce CH.sub.4 and C.sub.2+ paraffins for
recycling into the OCM reactor. A range of catalysts, including any
suitable F-T catalyst, can be employed. Reactor designs, including
those discussed in the present disclosure, can be employed. F-T
reactor operation conditions, including temperature and pressure,
can be optimized. This approach can reduce H.sub.2 consumption
compared to a methanation reactor.
Hydrocarbon Separations
In natural gas processing plants, methane can be separated from
ethane and higher carbon-content hydrocarbons (conventionally
called natural gas liquids or NGLs) to produce a methane-rich
stream that can meet the specifications of pipelines and sales gas.
Such separation can be performed using cryogenic separation, such
as with the aid of one or more cryogenic units.
The raw natural gas fed to gas processing plants can have a molar
composition of 70% to 95% methane and 4% to 20% NGLs, the balance
being inert gas(ses) (e.g., CO.sub.2 and N.sub.2). The ratio of
methane to ethane can be in the range of 5-25:1. Given the
relatively large amount of methane present in the stream fed to
cryogenic sections of gas processing plants, at least some or
substantially all of the cooling duty required for the separation
is provided by a variety of compression and expansion steps
performed on the feed stream and the methane product stream. None
or a limited portion of the cooling duty can be supplied by
external refrigeration units.
There are various approaches for separating higher carbon alkanes
(e.g., ethane) from natural gas, such as recycle split vapor (RSV)
and gas sub-cooled process (GSP) processes, which can maximize the
recovery of ethane (e.g., >95% recovery) while providing most or
all of the cooling duty via internal compression and expansion of
the methane itself. However, the application of such approach in
separating alkenes (e.g., ethylene) from an OCM product stream
comprising methane may result in a limited recovery (e.g., provide
less than 95% recovery) of the alkene product, due at least in part
to (i) the different vapor pressure of alkenes and alkanes, and/or
(ii) the presence of significant amounts of H.sub.2 in the OCM
product stream, which can change the boiling curve and,
particularly, the Joule-Thomson coefficient of the methane stream
that needs to be compressed and expanded to provide the cooling
duty. Hydrogen can display a negative or substantially low
Joule-Thomson coefficient, which can cause a temperature increase
or a substantially low temperature decrease in temperature when a
hydrogen-reach stream is expanded.
In some embodiments, the design of a cryogenic separation system of
an OCM-based plant can feature a different combination of
compression/expansion steps for internal refrigeration and, in some
cases, external refrigeration. The present disclosure provides a
separation system comprising one or more cryogenic separation units
and one or more de-methanizer units. Such a system can maximize
alkene recovery (e.g., provide greater than 95% recovery) from a
stream comprising a mixture of alkanes, alkenes, and other gases
(e.g., H.sub.2), such as in an OCM product stream (see FIGS. 24 and
25 and the associated text).
In such separation system, the cooling duty can be supplied by a
combination of expansion of the OCM effluent (feed stream to the
cryogenic section) when the OCM effluent pressure is higher than a
de-methanizer column; expansion of at least a portion or all of the
de-methanizer overhead methane-rich stream; compression and
expansion of a portion of the de-methanizer overhead methane-rich
stream; and/or external propane, propylene or ethylene
refrigeration units.
FIGS. 28-33 show various separation systems, as can be employed
with various systems and methods of the present disclosure. Such
systems can be employed for use in the OCM-ETL systems described
herein, such as used as the vapor-liquid separator 320 described
above in the context of FIGS. 3 and 4.
FIG. 28 shows a separation system 2800 comprising a first heat
exchanger 2801, a second heat exchanger 2802, a de-methanizer 2803,
and a third heat exchanger 2804. The direction of fluid flow is
shown in the figure. The de-methanizer 2803 can be a distillation
unit or multiple distillation units (e.g., in series). In such a
case, the de-methanizer can include a reboiler and a condenser,
each of which can be a heat exchanger. An OCM effluent stream 2805
is directed to the first heat exchanger 2801 at a pressure from
about 10 to 100 bar (absolute), or 20 to 40 bar. The OCM effluent
stream 2805 can include methane and C.sub.2+ compounds, and may be
provided in an OCM product stream from an OCM reactor (not shown).
The OCM effluent stream 2805 is then directed from the first heat
exchanger 2801 to the second heat exchanger 2802. In the first heat
exchanger 2801 and the second heat exchanger 2802, the OCM effluent
stream 2805 is cooled upon heat transfer to a de-methanizer
overhead stream 2806, a de-methanizer reboiler stream 2807, a
de-methanizer bottom product stream 2808, and a refrigeration
stream 2809 having a heat exchange fluid comprising propane or an
equivalent cooling medium, such as, but not limited to, propylene
or a mixture of propane and propylene.
The cooled OCM effluent 2805 can be directed to the de-methanizer
2803, where light components, such as CH.sub.4, H.sub.2 and CO, are
separated from heavier components, such as ethane, ethylene,
propane, propylene and any other less volatile component present in
the OCM effluent stream 2805. The light components are directed out
of the de-methanizer along the overhead stream 2806. The heavier
components are directed out of the de-methanizer along the bottom
product stream 2808. The de-methanizer can be designed such that at
least about 60%, 70%, 80%, 90%, or 95% of the ethylene in the OCM
effluent stream 2805 is directed to the bottom product stream
2808.
The de-methanizer overhead stream 2806 can contain at least 60%,
65%, or 70% methane. The overhead stream 2806 can be expanded
(e.g., in a turbo-expander or similar machine or flashed over a
valve or similar device) to decrease the temperature of the
overhead stream 2806 prior to directing the overhead stream 2806 to
the second heat exchanger 2802 and subsequently the first heat
exchanger 2801. The overhead stream 2806 can be cooled in the third
heat exchanger 2804, which can be cooled using a reflux stream and
a hydrocarbon-containing cooling fluid, such as, for example,
ethylene.
The overhead stream 2806, which can include methane, can be
recycled to an OCM reactor and/or directed for other uses, such as
a natural gas pipeline. In some examples, the bottom product
stream, which can contain C.sub.2+ compounds (e.g., ethylene), can
be directed to an ETL system.
FIG. 29 shows another separation system 2900 that may be employed
for use with systems and methods of the present disclosure. The
direction of fluid flow is shown in the figure. The system 2900
comprises a first heat exchanger 2901, de-methanizer 2902 and a
second heat exchanger 2903. The de-methanizer 2902 can be a
distillation unit or multiple distillation units (e.g., in series).
An OCM effluent stream 2904 is directed into the first heat
exchanger 2901. The OCM effluent stream 2904 can include methane
and C.sub.2+ compounds, and may be provided in an OCM product
stream from an OCM reactor (not shown). The OCM effluent stream
2904 can be provided at a pressure from about 10 bar (absolute) to
100 bar, or 40 bar to 70 bar. The OCM effluent stream 2904 can be
cooled upon heat transfer to a de-methanizer overhead streams 2905
and 2906 from the second heat exchanger 2903, a de-methanizer
reboiler stream 2907, and a refrigeration stream having a cooling
fluid comprising, for example, propane or an equivalent cooling
medium, such as, but not limited to, propylene or a mixture of
propane and propylene. In some cases, the de-methanizer overhead
streams 2905 and 2906 are combined into an output stream 2912
before or after passing through the first heat exchanger 2901.
Subsequent to cooling in the first heat exchanger 2901, the OCM
effluent stream 2904 can be expanded in a turbo-expander or similar
device or flashed over a valve or similar device to a pressure of
at least about 5 bar, 6 bar, 7 bar, 8 bar, 9 bar, or 10 bar. The
cooled OCM effluent stream 2904 can then be directed to the
de-methanizer 2902, where light components (e.g., CH.sub.4, H.sub.2
and CO) are separated from heavier components (e.g., ethane,
ethylene, propane, propylene and any other less volatile component
present in the OCM effluent stream 2904). The light components are
directed to an overhead stream 2909 while the heavier components
(e.g., C.sub.2+) are directed along a bottoms stream 2910. A
portion of the overhead stream 2909 is directed to second heat
exchanger 2903 and subsequently to the first heat exchanger 2901
along stream 2906. A remainder of the overhead stream 2909 is
pressurized in a compressor and directed to the second heat
exchanger 2903. The remainder of the overhead stream 2909 is then
directed to a phase separation unit 2911 (e.g., distillation unit
or vapor-liquid separator). Liquids from the phase separation unit
2911 are directed to the second heat exchanger 2903 and
subsequently returned to the de-methanizer 2902. Vapors from the
phase separation unit 2911 are expanded (e.g., in a turbo-expander
or similar device) and directed to the second heat exchanger 2903,
and thereafter to the first heat exchanger along stream 2905. The
de-methanizer 2902 can be designed such that at least about 60%,
70%, 80%, 90%, or 95% of the ethylene in the OCM effluent stream
2904 is directed to the bottom product stream 2910.
FIG. 30 shows another separation system 3000 that may be employed
for use with systems and methods of the present disclosure. The
direction of fluid flow is shown in the figure. The system 3000
comprises a first heat exchanger 3001, a de-methanizer 3002, a
second heat exchanger 3003 and a third heat exchanger 3004. The
system 3000 may not require any external refrigeration. The
de-methanizer 3002 can be a distillation unit or multiple
distillation units (e.g., in series). An OCM effluent stream 3005
is directed to the first heat exchanger 3001 at a pressure from
about 10 bar (absolute) to 100 bar, or 40 bar to 70 bar. In the
first heat exchanger 3001, the OCM effluent stream 3005 can be
cooled upon heat transfer to de-methanizer overhead streams 3006
and 3007, a de-methanizer reboiler stream 3008 and a de-methanizer
bottom product stream 3009. In some cases, the de-methanizer
overhead streams 3006 and 3007 are combined into a common stream
3015 before or after they are passed through the first heat
exchanger 3001. The OCM effluent stream 3005 is then expanded to a
pressure of at least about 5 bar, 6 bar, 7 bar, 8 bar, 9 bar, or 10
bar, such as, for example, in a turbo-expander or similar machine
or flashed over a valve or similar device. The cooled OCM effluent
stream 3005 is then directed to the de-methanizer 3002, where light
components (e.g., CH.sub.4, H.sub.2 and CO) are separated from
heavier components (e.g., ethane, ethylene, propane, propylene and
any other less volatile component present in the OCM effluent
stream 3005). The light components are directed to an overhead
stream 3010 while the heavier components are directed along the
bottom product stream 3009. The de-methanizer 3002 can be designed
such that at least about 60%, 70%, 80%, 90%, or 95% of the ethylene
in the OCM effluent stream 3005 is directed to the bottom product
stream 3009.
The de-methanizer overhead stream 3010, which can contain at least
50%, 60%, or 70% methane, can be divided into two streams. A first
stream 3011 is compressed in compressor 3012 and cooled in the
second heat exchanger 3003 and phase separated in a phase
separation unit 3013 (e.g., vapor-liquid separator or distillation
column). Vapors from the phase separation unit 3013 are expanded
(e.g., in a turbo-expander or similar device) to provide part of
the cooling duty required in heat exchangers 3001, 3003 and 3004.
Liquids from the phase separation unit 3013 are sub-cooled in the
third heat exchanger 3004 and recycled to the de-methanizer 3002. A
second stream 3014 from the overhead stream 3010 can be expanded
(e.g., in a turbo-expander or similar device) to decrease its
temperature and provide additional cooling to the heat exchangers
3001, 3003 and 3004.
FIG. 31 shows another separation system 3100 that may be employed
for use with systems and methods of the present disclosure. The
direction of fluid flow is shown in the figure. The system 3100
comprises a first heat exchanger 3101, a de-methanizer 3102, and a
second heat exchanger 3103. An OCM effluent stream 3104 is directed
to the first heat exchanger 3101 at a pressure from about 2 bar
(absolute) to 100 bar, or 3 bar to 10 bar. The first heat exchanger
3101 can interface with a propane refrigeration unit 3115 and/or an
ethylene refrigeration unit 3116. In the first heat exchanger 3101,
the OCM effluent stream 3104 can be cooled upon heat transfer to
de-methanizer overhead streams 3105 and 3106, a de-methanizer
reboiler stream, a de-methanizer pump-around stream, and various
levels of external refrigeration, such as using cooling fluids
comprising ethylene and propylene. In some cases, the de-methanizer
overhead streams 3105 and 3106 are combined into a single stream
3114 before or after they are cooled. The cooled OCM effluent
stream 3104 is then directed to the de-methanizer 3102, where light
components (e.g., CH.sub.4, H.sub.2 and CO) are separated from
heavier components (e.g., ethane, ethylene, propane, propylene and
any other less volatile component present in the OCM effluent
stream 3104). The light components are directed to an overhead
stream 3107 and the heavier components are directed along a bottom
product stream 3108. The de-methanizer 3102 can be designed such
that at least about 60%, 70%, 80%, 90%, or 95% of the ethylene in
the OCM effluent stream 3104 is directed to the bottom product
stream 3108.
The de-methanizer overhead stream, which can contain at least about
50%, 60%, or 70% methane, can be divided into two streams. A first
stream 3113 can be compressed in a compressor 3109, cooled in the
second heat exchanger 3103 and phase-separated in a phase
separation unit 3110 (e.g., distillation column or vapor-liquid
separator). Vapors from the phase separation unit 3110 can be
expanded (e.g., in a turbo-expander or similar device) to provide
part of the cooling duty required for the heat exchanger 3101 and
3103. Liquids from the phase separation unit 3110 can be sub-cooled
and flashed (e.g., over a valve or similar device), and the
resulting two-phase stream is separated in an additional phase
separation unit 3111. Liquids from the additional phase separation
unit 3111 are recycled to the de-methanizer 3102 and vapors from
the additional phase separation unit are mixed with expanded vapors
from the phase separation unit 3110 prior to being directed to the
second heat exchanger 3103.
A second stream 3112 from the overhead stream 3107 can be expanded
(e.g., in a turbo-expander or similar device) to decrease its
temperature and provide additional cooling for the heat exchanger
3101 and 3103. Any additional cooling that may be required for the
second heat exchanger 3103 can be provided by an external
refrigeration system, which may employ a cooling fluid comprising
ethylene or an equivalent cooling medium.
In some cases, recycle split vapor (RSV) separation can be
performed in combination with de-methanization.
In some instances, the methane undergoes an OCM and/or ETL process
to produce liquid fuel or aromatic compounds (e.g., higher
hydrocarbon liquids) and contains molecules that have gone through
methanation. In some embodiments, the compounds have been through a
recycle split vapor (RSV) separation process. In some cases,
alkanes (e.g., ethane, propane, butane) are cracked in a post-bed
cracker.
Systems above and elsewhere herein are not limited to ethylene and
may be configured to operate with other olefins, such as propylene,
butenes, pentene, or other alkenes. Although various systems and
methods herein have been described in the context of ethylene to
liquids, it will be appreciated that other alkenes may be used. For
example, an OCM reactor may generate an OCM effluent stream
comprising propylene and/or one or more butenes, which may be used
to provide one or more streams comprising higher molecular weight
hydrocarbons.
Systems of the present disclosure may be suitable for generating
liquids at less than or equal to about 250 kilotons per annum (KTA)
("small scale"), or generating liquids at greater than about 250
KTA ("world scale"). In some examples, a world scale OCM-ETL system
generates at least about 1000, 1100, 1200, 1300, 1400, 1500, or
1600 KTA of liquids.
Ethane Skimmers
The systems and methods described herein can process natural gas
into gas that is suitable for sale (i.e., "sales gas" that meets
the specifications required for transportation by pipeline). In
some cases, the systems and methods of the present disclosure can
convert methane and/or ethane (e.g., from natural gas) to sales gas
as well as products such as LPG, gasoline, distillate fuels, and/or
aromatic chemicals. Such a system or method is referred to as an
"ethane skimmer".
Ethane can be fed directly into a post-bed cracker (PBC), which can
be a portion of an OCM reactor downstream of the OCM catalyst,
where the heat generated in the OCM reaction can be used to crack
the ethane to ethylene. As an alternative, the PBC can be a unit
that is separate from the OCM reactor and in some cases in thermal
communication with the OCM reactor. The ethane feed stream to the
OCM reactor can include (a) ethane recycled to the OCM reactor from
an OCM reactor effluent stream, which can be separated in at least
one downstream separation module and recycled to the OCM reactor,
(b) ethane present in other feed streams (e.g., natural gas), which
can be separated in at least one separation module and recycled to
the OCM reactor, and (c) any additional (i.e., fresh) ethane
feed.
The maximum amount of ethane that can be converted in the PBC can
be limited by the flow rate of material exiting the OCM catalyst
and/or its temperature. It can be advantageous to utilize a high
proportion of the maximum amount of PBC. In some cases, the amount
of ethane converted to ethylene is about 50%, about 60%, about 70%,
about 80%, about 85%, about 90%, about 95%, or about 99% of the
maximum amount of ethane that can be converted to ethylene in the
PBC. In some instances, the amount of ethane converted to ethylene
is at least about 50%, at least about 60%, at least about 70%, at
least about 80%, at least about 85%, at least about 90%, at least
about 95%, or at least about 99% of the maximum amount of ethane
that can be converted to ethylene in the PBC.
Achieving a high proportion of the maximum PBC capacity can be
accomplished by adding natural gas to the system, which can have a
concentration of ethane that depends on many factors, including the
geography and type and age of the natural gas well. The treatment
and separation modules of the process described herein can be used
to purify or fractionate the ETL effluent, and can additionally be
used to treat (e.g., remove water and CO.sub.2) and purify the
natural gas that is added to the system along with the ETL
effluent, such as, e.g., by separating C.sub.2+ compounds from
methane and separating ethane from ethylene. In some cases, ethane
contained in the natural gas feed can be recycled to the OCM
reactor (e.g., PBC region) as pure ethane and the system may not be
sensitive to the purity and composition of the natural gas, making
raw natural gas a suitable input to the system.
The maximal PBC capacity can depend on the ratio between methane
and ethane in the input to the OCM reactor, including in some
instances the PBC portion. In some cases, the PBC capacity is
saturated when the molar ratio of methane to ethane is about 1,
about 2, about 3, about 4, about 5, about 6, about 7, about 8,
about 9, about 10, about 11, about 12, about 13, about 14, or about
15. In some cases, the PBC capacity is saturated when the molar
ratio of methane to ethane is at least about 1, at least about 2,
at least about 3, at least about 4, at least about 5, at least
about 6, at least about 7, at least about 8, at least about 9, at
least about 10, at least about 11, at least about 12, at least
about 13, at least about 14, or at least about 15. In some cases,
the PBC capacity is saturated when the molar ratio of methane to
ethane is at most about 5, at most about 6, at most about 7, at
most about 8, at most about 9, at most about 10, at most about 11,
at most about 12, at most about 13, at most about 14 or at most
about 15. In some cases, the PBC capacity is saturated when the
molar ratio of methane to ethane is between about 7 and 10 parts
methane to one part ethane.
Natural gas (raw gas or sales gas) can have a concentration of
ethane of less than about 30 mol %, 25 mol %, 20 mol %, 15 mol %,
10 mol %, 9 mol %, 8 mol %, 7 mol %, 6 mol %, 5 mol %, 4 mol %, 3
mol %, 2 mol % or 1 mol %. In some cases, natural gas has a methane
to ethane ratio greater than about 1:1, 2:1, 3:1, 4:1, 5:1, 6:1,
7:1, 8:1, 9:1, 10:1, 11:1, 12:1, 13:1, 14:1, 15:1, 16:1, 17:1,
18:1, 19:1, 20:1 or 40:1. The ethane skimmer implementation
described herein can be used to inject more natural gas feed into
the system than what may be required to produce the desired or
predetermined amount of ethylene or other products. The excess
methane can be drawn from a stream downstream of the methanation
unit and sold as sales gas (which may lack an appreciable amount of
ethane but can still meet pipeline specifications and/or can be
directed to a power plant for power production). The ethane in the
additional natural gas feed can be used to saturate the PBC
capacity. Any excess ethane can be drawn from the C.sub.2 splitter
and exported as pure ethane. The ethane skimmer implementation
described herein can result in additional product streams from the
system (namely sales gas, natural gas liquids, gasoline, diesel or
jet fuels and/or aromatic chemicals). In such a case, the process
can be used to achieve both natural gas processing and production
of C.sub.2+ chemicals or fuels.
The ethane skimmer implementation can be readily understood by
reference to FIG. 32. Natural gas 3200 can be fed into a
desulfurization unit 3202 and then into a gas compressor 3204.
Oxygen can be provided from an air separation unit 3206 that can be
powered by a gas turbine and combination cycle 3208 that is powered
by combustion of a portion of the natural gas and/or methane. The
oxygen and methane 3210 produced by the process can be injected
into an OCM reactor 3212 having a PBC portion 3214. The OCM
effluent can be fed into the process gas compressor 3204 and then
into the ETL module 3216. Products of the ETL module can be dried
in a drier 3218. The separation module can comprise a de-methanizer
3220, a de-ethanizer 3222 and a de-butanizer 3224. The
de-methanizer can separate C.sub.1 compounds from C.sub.2+
compounds and direct the C.sub.1 compounds (e.g., methane, carbon
monoxide and carbon dioxide) to a methanator 3226. The C.sub.1
compound stream can have any amount of C.sub.2+ compounds (e.g.,
about 0.5%, about 1%, about 1.5%, about 2%, about 2.5%, about 3%,
or about 3.5%). The methanator can convert the carbon monoxide
and/or carbon dioxide to methane (e.g., using hydrogen generated in
the process). The methane can be divided into any number of streams
that can be directed to the OCM reactor 3212, the gas turbine 3208,
and/or a pipeline 3228 or other means for delivering a methane
product to the market (i.e., sales gas). The ethane 3230 from the
separation module can be directed to the PBC. The system can
produce C.sub.3+ products such as liquified petroleum gas (LPG;
having C.sub.3 and C.sub.4 molecules) 3232 and C.sub.5+ products
3234 such as gasoline, diesel fuel, jet fuel, and/or aromatic
chemicals.
Overall, in the ethane skimmer process as shown in FIG. 32, at
least some or most (e.g., >70%, >80%, >85%, >90%,
>95%, or >99%) of the methane in the natural gas feed 3200
ends up in the methane recycle 3210, at least some or most (e.g.,
>70%, >80%, >85%, >90%, >95%, or >99%) of the
ethane in the natural gas feed ends up in the ethane recycle stream
3230, at least some or most (e.g., >70%, >80%, >85%,
>90%, >95%, or >99%) propane in the natural gas feed ends
up in the C.sub.3+ products streams 3232 and 3234. In some cases,
and ethane is added (not shown in FIG. 32) up to the point where
the PBC cracking capacity is saturated or nearly saturated (e.g.,
>70%, >80%, >85%, >90%, >95%, or >99%). Excess
ethane (e.g., beyond what is needed to saturate the PBC) can end up
in an ethane product stream (not shown). The ethane skimmer
implementation does not require a separate (i.e., fresh) ethane
stream to saturate or nearly saturate the PBC capacity of the
system.
Additional Products and Processes
In addition to the ethylene conversion processes described herein,
components other than ethylene that are produced in an ethylene
production process, e.g., contained within an OCM effluent gas, may
be directed to, and thus fluidly connected to additional conversion
processes. In particular, the OCM reaction process generates a
number of additional products, other than ethylene, including for
example, hydrogen gas (H.sub.2) and carbon monoxide (CO). In some
cases, the H.sub.2 and CO components of the OCM reaction product
slate are subjected to additional processing to produce other
products and intermediates, e.g., dimethylether (DME), methanol,
and hydrocarbons. These components may be useful in a variety of
different end products, including liquid fuels, lubricants and
propellants. In some embodiments, the H.sub.2 and CO components of
the OCM reaction effluent are separated from the other OCM
products. The H.sub.2 and CO can then be subjected to any of a
variety of syngas-like conversion processes to produce a variety of
different products, e.g., methanol, dimethylether, hydrocarbons,
lubricants, waxes and fuels or fuel blendstocks. In one example,
the H.sub.2 and CO components are subjected to a catalytic process
to produce DME via a methanol intermediate. The catalytic process
is described in detail in, e.g., U.S. Pat. No. 4,481,305, the full
disclosure of which is incorporated herein by reference in its
entirety for all purposes.
As noted herein, the ethylene conversion processes employed in the
integrated processes and systems of the invention may produce
olefinic products for use in a variety of different end products or
applications. For example, a portion or all of the ethylene
produced by the OCM process may be routed through one or more
catalytic processes or systems to oligomerize ethylene into LAOs of
ranging carbon numbers. These compounds can be particularly useful
in chemical manufacturing, e.g., in the production of amines, amine
oxides, oxo-alcohols, alkylated aromatics epoxides, tanning oils,
synthetic lubricants, lubricant additives, alpha olefin sulfonates,
mercaptans, organic alkyl aluminum, hydrogenated oligomers, and
synthetic fatty acids. Alternatively or additionally, the ethylene
may be oligomerized through LAO processes to produce
C.sub.4-C.sub.20 LAOs for use as liquid blend stocks for gasoline,
diesel or jet fuels. These LAOs can also be hydrogenated to linear
alkanes for fuel blend stocks for gasoline, jet, and diesel
fuel.
Processes used for the production of product ranges, e.g.,
C.sub.4-C.sub.30 LAOS, are generally referred to herein as "full
range processes" or "narrow range processes", as they produce a
range of chemical species, e.g., LAOs of varying chain length such
as 1-butene, 1-hexene, 1-octene, 1-decene, etc., in a single
process. Products from full range or narrow range processes may be
distilled or fractionated into, e.g., C.sub.4-C.sub.10 LAOs for use
as chemical process feedstocks, C.sub.10-C.sub.20 LAOs for use as a
jet fuel blendstock, diesel fuel blendstock, and chemical
feedstock. By contrast, processes that produce a single product
species in high yield, e.g., LAO of a single chain length such as
1-butene, 1-hexene, 1-octene, 1-decene or the like, are referred to
generally as selective processes.
Full and narrow ranges of products may be prepared from ethylene
using a variety of LAO processes, such as, for example, the
.alpha.-Sablin.RTM. process (See, e.g., Published International
Patent Application No. WO 2009/074203, European Patent No. EP
1749806B1, and U.S. Pat. No. 8,269,055, the full disclosures of
which are incorporated herein by reference in their entirety for
all purposes), the Shell higher olefin process (SHOP), the
Alphabutol process, the Alphahexol process, the AlphaSelect
process, the Alpha-Octol process, Linear-1 process, the Linealene
process, the Ethyl Process, the Gulftene process, and the Phillips
1-hexene process.
Briefly, the .alpha.-Sablin process employs a two-component
catalyst system of a zirconium salt and an aluminum alkyl
co-catalyst, for homogenous, liquid phase oligomerization of
ethylene to a narrow range of LAOS. The catalytic cycle comprises a
chain growth step by an ethylene insertion reaction at the
co-ordination site and displacement of the coordinated hydrocarbon
from the organometallic complex. The ratio of zirconium to aluminum
can be used to adjust between chain growth and displacement,
thereby adjusting the product spectrum more toward lighter or
heavier LAOS. For example, with a high Zr:Al ratio, the product
spectrum can be shifted to upwards of 80% C4-C8 LAOS, while lower
Zr:Al ratios will shift the product spectrum towards heavier LAOS.
The reaction is generally carried out in a bubble column reactor
with a solvent, such as toluene, and catalyst being fed into the
liquid phase at temperatures of between about 60.degree. C. and
100.degree. C. and pressures of between about 20 bar and 30 bar.
The liquid LAOs are then sent to a separation train to deactivate
the catalyst, separate the solvent and optionally perform any
additional product separations that are desired.
Additionally, all or a portion of these olefinic products may be
hydrogenated prior to distillation to convert the olefins into the
corresponding alkanes for use as alkane blendstocks for fuel
products, and then again, subjected to a distillation or other
separation process to produce the desired products.
In various embodiments, a wide range of other ethylene conversion
processes may be integrated at the back end of the OCM processes
described above, depending upon the desired product or products for
the overall process and system. For example, in alternative or
additional aspects, an integrated ethylene conversion process for
production of LAOs may include the SHOP system, a full range
ethylene conversion process which may be used to produce LAOs in
the C.sub.6-C.sub.16 range. Briefly, the SHOP system employs a
nickel-phosphine complex catalyst to oligomerize ethylene at
temperatures of from about 80.degree. C. to about 120.degree. C.,
and pressures of from about 70 bar to about 140 bar.
A variety of other full-range ethylene conversion processes may be
employed, including without limitation, the AlphaSelect process,
the Alpha-Octol process, Linear-1 process, the Linealene process,
the Synthol process, the Ethyl Process, the Gulftene process, the
Phillips 1-hexene process, and others. These processes are well
characterized in the literature, and reported, for example at the
Nexant/Chemsystems PERP report, Alpha Olefins, January 2004, the
full disclosure of which are incorporated herein by reference in
their entirety for all purposes.
As an alternative or in addition to full and/or narrow range
ethylene conversion processes, ethylene conversion processes that
may be integrated into the overall systems of the invention include
processes for the selective production of high purity single
compound LAO compositions. As used herein, processes that are
highly selective for the production of a single chemical species
are generally referred to as selective or "on purpose" processes,
as they are directed at production of a single chemical species in
high selectivity. In the context of LAO production, such on purpose
processes will typically produce a single LAO species, e.g.,
1-butene, 1-hexene, 1-octene, etc., at selectivities of greater
than 50%, in some cases greater than 60%, greater than 75%, and
even greater than 90% selectivity for the single LAO species.
Examples of such on purpose processes for ethylene conversion to
LAOs include, for example, the Alphahexol process from IFP, the
Alphabutol process, or the Phillips 1-hexene process for the
oligomerization of ethylene to high purity 1-hexene, as well as a
wide range of other processes that may be integrated with the
overall OCM reactor system.
The Alphahexol process, for example, is carried out using phenoxide
ligand processes. In particular, ethylene trimerization may be
carried out using a catalytic system that involves a chromium
precursor, a phenoxyaluminum compound or alkaline earth phenoxide
and a trialkylaluminum activator at 120.degree. C. and 50 bar
ethylene pressure (See, e.g., U.S. Pat. No. 6,031,145, and European
Patent No. EP1110930, the full disclosures of which are
incorporated herein by reference in their entirety for all
purposes). Likewise, the Phillips 1-hexene process employs a
chromium(III) alkanoate, such as chromium tris(2-ethylhexanoate,
pyrrole, such as 2,5-dimethylpyrrole, and Et3Al to produce 1-hexene
at high selectivity, e.g., in excess of 93%. See, e.g., European
Patent No. EP0608447 and U.S. Pat. No. 5,856,257, the full
disclosures of which are incorporated herein by reference in their
entirety for all purposes. A variety of other ethylene
trimerization processes may be similarly integrated to the back end
of the OCM systems described herein. These include, for example,
the British Petroleum PNP trimerization system (see, e.g.,
Published International Patent Application No. WO 2002/04119, and
Carter et al., Chem. Commun. 2002, 858), and Sasol PNP
trimerization system (see, e.g., Published International Patent
Application No. WO2004/056479, discussed in greater detail), the
full disclosures of which are incorporated herein by reference in
their entirety for all purposes.
The Alphabutol process employs a liquid phase proprietary soluble
catalyst system of Ti(IV)/AlEt3, in the dimerization of ethylene to
1-butene at relatively high purity, and is licensed through Axens
(Rueil-Malmaison, France). Ethylene is fed to a continuous liquid
phase dimerization reactor. A pump-around system removes the
exothermic heat of reaction from the reactor. The reactor operates
between 50-60.degree. C. at 300-400 psia. The catalyst is removed
from the product effluent and is ultimately fed to the 1-butene
purification column where comonomer-grade 1-butene is produced.
Still other selective ethylene conversion processes include the
catalytic tetramerization of ethylene to 1-octene. For example, one
exemplary tetramerization process employs a liquid phase catalytic
system using a Cr(III) precursor, such as [Cr(acac)3] or
[CrCl3(THF)3] in conjunction with a bis(phosphine)amine ligand and
a methylaluminooxane (MAO) activator at temperatures of between
about 40.degree. C. and 80.degree. C. and ethylene pressures of
from 20 to 100 bar, to produce 1-octene with high selectivity. See,
e.g., Published International Patent Application No. WO2004/056479
and Bollmann, et al., "Ethylene Tetramerization: A New Route to
Produce 1-Octene in Exceptionally High Selectivities" J. Am. Chem.
Soc., 2004, 126 (45), pp 14712-14713, the full disclosures of which
are incorporated herein by reference in their entirety for all
purposes.
In addition to the LAO processes described herein, ethylene
produced from the integrated OCM reactor systems can also be used
to make olefinic non-LAO linear hydrocarbons and branched olefinic
hydrocarbons through the same or different integrated processes and
systems. For example, the ethylene product from the OCM reactor
system may be passed through integrated reactor systems configured
to carry out the SHOP process, the Alphabutol process, the
Alphahexol process, the AlphaSelect process, the Alpha-Octol
process, Linear-1 process, the Linealene process, the Ethyl
Process, the Gulftene process, and/or the Phillips 1-hexene
process, to yield the resultant LAO products. The output of these
systems and processes may then be subjected to an olefin
isomerization step to yield linear olefins other than LAOS,
branched olefinic hydrocarbons, or the like. In addition, olefinic
non-LAO linear hydrocarbons and branched olefinic hydrocarbons can
be prepared by ethylene oligomerization over heterogeneous
catalysts such as zeolites, amorphous silica/alumina, solid
phosphoric acid catalysts, as well as doped versions of the
foregoing catalysts.
Other oligomerization processes have been described in the art,
including the olefin oligomerization processes set forth in
Published U.S. Patent Application No. 2012/0197053 (incorporated
herein by reference in its entirety for all purposes), which
describes processes used for production of liquid fuel components
from olefinic materials.
Although a number of processes are described with certain
specificity, that description is by way of example and not
limitation. In particular, it is envisioned that the full range of
ethylene oligomerization and/or conversion processes may be readily
integrated onto the back end of the OCM reactor systems for
conversion of methane to ethylene product, and subsequently to a
wide range of different higher hydrocarbon products. As noted
previously, certain embodiments of the ethylene conversion
processes that are integrated into the overall systems of the
invention are those that yield liquid hydrocarbon products. Other
embodiments of the ethylene conversion processes that are
integrated in the overall systems include process that are
particularly well-suited for use with dilute ethylene feed stocks
which optionally comprise additional components such as higher
hydrocarbons, unreacted OCM starting material (methane and/or other
natural gas components) and/or side products of the OCM reactions.
Examples of such other components are provided herein.
In addition to or as an alternative, the ethylene product produced
from the OCM reactor system may be routed through one or more
catalytic or other systems and processes to make non-olefinic
hydrocarbon products. For example, saturated linear and branched
hydrocarbon products may be produced from the ethylene product of
the OCM reactor system through the hydrogenation of the products of
the olefinic processes described above, e.g., the SHOP process, the
Alphabutol process, the Alphahexol process, the AlphaSelect
process, the Alpha-Octol process, Linear-1 process, the Linealene
process, the Ethyl Process, the Gulftene process, and/or the
Phillips 1-hexene process.
Other catalytic ethylene conversions systems that may be employed
include reacting ethylene over heterogeneous catalysts, such as
zeolites, amorphous silica/alumina, solid phosphoric acid
catalysts, and/or doped forms of these catalysts, to produce
mixtures of hydrocarbons, such as saturated linear and/or branched
hydrocarbons, saturated olefinic cyclic hydrocarbons, and/or
hydrocarbon aromatics. By varying the catalysts and or the process
conditions, selectivity of the processes for specific components
may be enhanced. For example, ethylene purified from OCM effluent
or unpurified OCM effluent containing ethylene can be flowed across
a zeolite catalyst, such as ZSM-5, or amorphous silica/alumina
material with SiO.sub.2/Al.sub.2O.sub.3 ratios of 23-280, at
ethylene partial pressures between 0.01 bar to 100 bar (undoped, or
doped with Zn and/or Ga in some embodiments or some combination
thereof) at temperatures above 350.degree. C. to give high liquid
hydrocarbon yield (80+%) and high aromatic selectivity (benzene,
toluene, xylene (BTX) selectivity>90% within the liquid
hydrocarbon fraction). Ethylene purified from OCM effluent or
unpurified OCM effluent containing ethylene can be flowed across a
zeolite catalyst, such as ZSM-5, or amorphous silica/alumina
material with SiO.sub.2/Al.sub.2O.sub.3 ratios of 23-280, at
ethylene partial pressures between 0.01 bar to 100 bar (undoped, or
with dopants including but not limited to, e.g., Ni, Mg, Mn, Ca,
and Co, or some combination of these) at temperatures above
200.degree. C., to give high liquid hydrocarbon yield (80+%) and
high gasoline selectivity (gasoline selectivity>90% within the
liquid hydrocarbon fraction). Ethylene purified from OCM effluent
or unpurified OCM effluent containing ethylene can be flowed across
a zeolite catalyst, such as ZSM-5, or amorphous silica/alumina
material with SiO.sub.2/Al.sub.2O.sub.3 ratios of 23-280 or a solid
phosphoric acid catalyst, at ethylene partial pressures between
0.01 bar to 100 bar at temperatures above 200.degree. C. to give
high liquid hydrocarbon yield (80+%) and high distillate
selectivity (gasoline selectivity>90% within the liquid
hydrocarbon fraction).
In some embodiments, to achieve high jet/diesel fuel yields, a two
oligomerization reactor system is used in series. The first
oligomerization reactor takes the ethylene and oligomerizes it to
C.sub.3-C.sub.6 olefins over modified ZSM-5 catalysts, e.g., Mg,
Ca, or Sr doped ZSM-5 catalysts. The C.sub.3-C.sub.6 olefins can be
the end products of the process or alternatively can be placed in a
second oligomerization reactor to be coupled into jet/diesel fuel
range liquid.
In addition, some embodiments of the ethylene conversion processes
also include processes for production of oxygenated hydrocarbons,
such as alcohols and/or epoxides. For example, the ethylene product
can be routed through an integrated system that includes a
heterogeneous catalyst system, such as a solid phosphoric acid
catalyst in the presence of water, to convert the ethylene to
ethanol. This process has been routinely used to produce 200 proof
ethanol in the process used by LyondellBasell. In other
embodiments, longer chain olefins and/or LAO's, derived from OCM
ethylene by oligomerization, can be likewise converted to alkyl
alcohols using this same process. See, e.g., U.S. Pat. Nos.
2,486,980; 3,459,678; 4,012,452, the full disclosures of which are
incorporated herein by reference in their entirety for all
purposes. In alternate embodiments, ethylene undergoes a vapor
oxidation reaction to make ethylene oxide over a silver based
catalyst at 200-300.degree. C. at 10-30 atmospheres of pressure
with high selectivity (80+%). Ethylene oxide is an important
precursor for synthesis of ethylene glycol, polyethylene glycol,
ethylene carbonate, ethanolamines, and halohydrins. See, e.g.,
Chemsystems PERP Report Ethylene Oxide/Ethylene Glycol 2005, which
is herein incorporated by reference.
In still other aspects, the ethylene product produced from the OCM
reactor system may be routed to a reactor system that reacts the
ethylene with various halogen sources (acids, gases, and others) to
make halogenated hydrocarbons useful, for example, as monomers in
producing halogenated polymers, such as polyvinyl chloride (PVC).
For example, in one ethylene dichloride (EDC) process, available
from ThyssenKrupp Uhde, ethylene can be reacted with chlorine gas
to make EDC, an important precursor to vinyl-chloride monomer (VCM)
for polyvinylchloride (PVC) production. This process also can be
modified EDC to react ethylene with hydrochloric acid (HCl) to make
EDC via oxychlorination.
In still other exemplary ethylene conversion processes, the
ethylene product of the OCM reactor system may be converted to
alkylated aromatic hydrocarbons, which are also useful as chemical
and fuel feedstocks. For example, in the Lummus CD-Tech EB process
and the Badger EB process, benzene can be reacted with OCM
ethylene, in the presence of a catalyst, to make ethylbenzene. See,
e.g., U.S. Pat. No. 4,107,224, the full disclosure of which is
incorporated herein by reference in its entirety for all purposes.
Ethylbenzene can be added to gasoline as a high-octane gasoline
blendstock or can be dehydrogenated to make styrene, the precursor
to polystyrene.
In addition to the liquid and other hydrocarbons described above,
in certain aspects, one or more of the integrated ethylene
conversion processes is used to convert ethylene product from the
OCM reactor system to one or more hydrocarbon polymers or polymer
precursors. For example, in some embodiments ethylene product from
the integrated OCM reactor systems is routed through an integrated
Innovene process system, available through Ineos Technologies,
Inc., where the ethylene is polymerized in the presence of a
catalyst, in either a slurry or gas phase system, to make long
hydrocarbon chains or polyethylene. By varying the process
conditions and catalyst the process and system can be used to
produce high density polyethylene or branched low density
polyethylene, etc. The Innovene G and Innovene S processes are
described at, for example, at "Ineostechnologies.com". See also
Nexant/Chemsystems HDPE Report, PERP 09/10-3, January 2011, the
full disclosure of which is incorporated herein by reference in its
entirety for all purposes.
Alternatively, ethylene from OCM can be introduced, under high
pressure, into an autoclave or tubular reactor in the presence of a
free radical initiator, such as O.sub.2 or peroxides, to initiate
polymerization for the preparation of low-density polyethylene
(LDPE). See e.g., "Advanced Polyethylene Technologies" Adv Polym
Sci (2004) 169:13-27, the full disclosure of which is incorporated
herein by reference in its entirety for all purposes.
Alternatively, ethylene from OCM can be introduced, under low
pressure in the presence of a chromium oxide based catalyst,
Ziegler-Natta catalyst, or a single-site (metallocene or non
metallocene based) catalyst, to prepare HDPE, MDPE, LLDPE, mLLDPE,
or bimodal polyethylene. The reactor configurations for synthesis
of HDPE, LLDPE, MDPE, and biomodal PE can be a slurry process, in
which ethylene is polymerized to form solid polymer particles
suspended in a hydrocarbon diluent, a solution process in which
dissolved ethylene is polymerized to form a polymer dissolved in
solvent, and/or a gas phase process in which ethylene is
polymerized to form a solid polymer in a fluidized bed of polymer
particles. Ethylene from OCM can be co-polymerized with different
monomers to prepare random and block co-polymers. Co-monomers for
ethylene copolymerization include but are not limited to: at least
one olefin comonomer having three to fifteen carbons per molecule
(examples are propylene and LAO's such as 1-butene, 1-hexene,
1-octene), oxygenated co-monomers such as: carbon oxide; vinyl
acetate, methyl acrylate; vinyl alcohols; allyl ethers; cyclic
monomers such as norbornene and derivatives thereof; aromatic
olefins such as: styrene and derivatives thereof. These ethylene or
LAO copolymerization processes, e.g., where ethylene is
copolymerized with different monomers, are generally referred to
herein as copolymerization processes or systems.
More exemplary ethylene conversion processes that may be integrated
with the OCM reactor systems include processes and systems for
carrying out olefin metathesis reactions, also known as
disproportionation, in the production of propylene. Olefin
metathesis is a reversible reaction between ethylene and butenes in
which double bonds are broken and then reformed to form propylene.
"Propylene Production via Metathesis, Technology Economics Program"
by Intratec, ISBN 978-0-615-61145-7, Q2 2012, the full disclosure
of which is incorporated herein by reference in its entirety for
all purposes. Propylene yields of about 90 wt % are achieved. This
option may also be used when there is no butene feedstock. In this
case, part of the ethylene from the OCM reaction feeds into an
ethylene-dimerization unit that converts ethylene into butene.
As noted herein, one, two, three, four or more different ethylene
conversion processes are provided integrated into the overall
systems of the invention, e.g., as shown in FIG. 1. As will be
appreciated, these ethylene conversion systems will include fluid
communications with the OCM systems described herein, and may be
within the same facility or within an adjacent facility. Further,
these fluid communications may be selective. In particular, in
certain embodiments the interconnect between the OCM system
component and the ethylene conversion system component(s) is able
to selectively direct all of an ethylene product from the OCM
system to any one ethylene conversion system at a given time, and
then direct all of the ethylene product to a second different
ethylene conversion system component at a different time.
Alternatively, such selective fluid communications may also
simultaneously direct portions of the ethylene product to two or
more different ethylene conversion systems to which the OCM system
is fluidly connected.
These fluid communications will typically comprise interconnected
piping and manifolds with associated valving, pumps, thermal
controls and the like, for the selective direction of the ethylene
product of the OCM system to the appropriate ethylene conversion
system component or components.
In an example, ethylene produced by the methods described herein
(e.g., by OCM) can be converted into 1-butene or 2-butene. In some
cases, ETL methods and systems provided herein can be used to form
1-butene but no appreciable 2-butene, or 2-butene but no
appreciable 1-butene. Methods for generating 1-butene from ethylene
are disclosed in U.S. Pat. No. 2,943,125, U.S. Pat. No. 3,686,350,
U.S. Pat. No. 4,101,600, U.S. Pat. No. 8,624,042, and U.S. Pat. No.
5,792,895, each of which is entirely incorporated herein by
reference.
As an alternative or in addition to, ethylene produced by the
methods described herein (e.g., by OCM) can be converted into
1-hexene. Methods for converting ethylene to 1-hexene are described
in U.S. Pat. No. 6,380,451, U.S. Pat. No. 7,157,612, U.S. Pat. No.
5,057,638, U.S. Pat. No. 8,658,750, and U.S. Pat. No. 5,811,618,
each of which is entirely incorporated herein by reference. As an
alternative or in addition to, ethylene produced by the methods
described herein (e.g., by OCM) can be converted into 1-octene.
Methods for converting ethylene to 1-Octene are described in U.S.
Pat. No. 5,292,979, U.S. Pat. No. 5,811,619, U.S. Pat. No.
5,817,905, and U.S. Pat. No. 6,103,654, each of which is entirely
incorporated herein by reference.
In some cases, ethylene produced by the methods described herein
(e.g., by OCM) can be converted into C4 to C18 and higher
.alpha.-olefins (1-butene, 1-hexene, 1-octene, 1-decene and
higher). Oligomerization of ethylene into linear alpha olefins
(LAO) can be carried out in a bubble column reactor with the
solvent and the dissolved catalyst components fed to the liquid
phase. Methods for converting ethylene to C.sub.4-C.sub.18 and
higher .alpha.-olefins are described in Canadian Patent Application
Number CA 2,765,769, German Patent Number DE 4338414, German Patent
Number DE 4338416, U.S. Pat. No. 3,862,257, U.S. Pat. No.
4,966,874, and U.S. Pat. No. 5,449,850, each of which is entirely
incorporated herein by reference.
As an alternative or in addition to, ethylene produced by the
methods described herein (e.g., by OCM) can be converted into
C.sub.4 to C.sub.10 .alpha.-olefins (1-butene, 1-hexene, 1-octene,
and 1-decene). Methods for converting ethylene to C.sub.4-C.sub.10
.alpha.-olefins are described in U.S. Pat. No. 3,660,519, U.S. Pat.
No. 3,584,071, European Patent Number EP 0,722,922, U.S. Pat. No.
4,314,090, U.S. Pat. No. 5,345,023, and U.S. Pat. No. 6,221,986,
each of which is entirely incorporated herein by reference.
In another example, ethylene produced by the methods described
herein (e.g., by OCM) can be converted into propylene (propene).
For example, n-butenes can be reacted with ethylene using a
heterogeneous catalyst system in a fixed bed reactor process.
Methods for converting ethylene to propylene are described in U.S.
Pat. No. 6,683,019, U.S. Pat. No. 7,214,841, U.S. Pat. No.
8,153,851, and U.S. Pat. No. 8,258,358, each of which is entirely
incorporated herein by reference.
As an alternative or in addition to, ethylene produced by the
methods described herein (e.g., by OCM) can be converted into
ethylene dichloride (EDC). For example, ethylene can be reacted
with chlorine in liquid phase in presence of a catalyst system.
Methods for converting ethylene to EDC are described in German
Patent Number DE 19 05 517, German Patent Number DE 25 40 257,
German Patent Number DE 40 39 960 A16, U.S. Pat. No. 7,579,509,
U.S. Pat. No. 7,671,244, and U.S. Pat. No. 6,841,708, each of which
is entirely incorporated herein by reference.
Ethylene produced by the methods described herein (e.g., by OCM)
can be converted into high density polyethylene (HDPE) or other
types of polyethylene. For example, ethylene or a mixture of
ethylene with one or more alpha olefins can be reacted in the gas
phase in the presence of a catalyst system. Methods for converting
ethylene to HDPE are described in U.S. Pat. No. 5,473,027, U.S.
Pat. No. 5,473,027, U.S. Pat. No. 6,891,001, and U.S. Pat. No.
4,882,400, each of which is entirely incorporated herein by
reference.
The ethylene produced by the methods described herein (e.g., by
OCM) can be converted into ethanol. For example, a mixture of
ethylene and water is reacted over a heterogeneous catalyst (e.g.,
solid phosphoric acid catalyst) in a reactor to form ethanol by
direct hydration of ethylene. Methods for converting ethylene to
ethanol are described in U.S. Pat. No. 2,486,980, U.S. Pat. No.
2,579,601, U.S. Pat. No. 2,673,221, and U.S. Pat. No. 3,686,334,
each of which is entirely incorporated herein by reference.
Acetylene can be selectively hydrogenated to ethylene while present
in a mixture containing ethylene and other components without
hydrogenating ethylene. For example, a feed containing acetylene
and ethylene is reacted in the presence of hydrogen over a
heterogeneous catalyst in a fixed bed reactor system. Methods for
selective hydrogenating acetylene are described in U.S. Pat. No.
3,128,317, U.S. Pat. No. 4,126,645, U.S. Pat. No. 4,367,353, U.S.
Pat. No. 4,329,530, U.S. Pat. No. 4,440,956, U.S. Pat. No.
5,414,170, U.S. Pat. No. 6,509,292, and Xu, Ling, et al. "Maximise
ethylene gain and acetylene selective hydrogenation efficiency,"
Petroleum technology quarterly 18.3 (2013): 39-42, each of which is
entirely incorporated herein by reference.
Acetylene and dienes, such as butadiene, can be selectively
hydrogenated while present in a mixture containing ethylene and
other components without hydrogenating the ethylene present. For
example, a feed containing acetylene and dienes is reacted in the
presence of hydrogen over a heterogeneous catalyst in a fixed bed
reactor system. Methods for selective hydrogenating acetylene and
dienes are described in U.S. Pat. No. 3,900,526, U.S. Pat. No.
5,679,241, U.S. Pat. No. 6,759,562, U.S. Pat. No. 5,877,363, U.S.
Pat. No. 7,838,710, and U.S. Pat. No. 8,227,650, each of which is
entirely incorporated herein by reference.
Olefin to Liquids Reactors
Control Systems
The present disclosure provides computer control systems that can
be employed to regulate or otherwise control the methods and
systems provided herein. A control system of the present disclosure
can be programmed to control process parameters to, for example,
effect a given product distribution, such as a higher concentration
of alkenes as compared to alkanes in a product stream out of an OCM
and/or ETL reactor.
FIG. 33 shows a computer system 3301 that is programmed or
otherwise configured to regulate OCM and/or ETL reactions, such as
regulate fluid properties (e.g., temperature, pressure and stream
flow rate(s)), mixing, heat exchange and OCM and/or ETL reactions.
The computer system 3301 can regulate, for example, fluid stream
("stream") flow rates, stream temperatures, stream pressures, OCM
and/or ETL reactor temperature, OCM and/or ETL reactor pressure,
the quantity of products that are recycled, and the quantity of a
first stream (e.g., methane stream) that is mixed with a second
stream (e.g., air stream).
The computer system 3301 includes a central processing unit (CPU,
also "processor" and "computer processor" herein) 3305, which can
be a single core or multi core processor, or a plurality of
processors for parallel processing. The computer system 3301 also
includes memory or memory location 3310 (e.g., random-access
memory, read-only memory, flash memory), electronic storage unit
3315 (e.g., hard disk), communication interface 3320 (e.g., network
adapter) for communicating with one or more other systems, and
peripheral devices 3325, such as cache, other memory, data storage
and/or electronic display adapters. The memory 3310, storage unit
3315, interface 3320 and peripheral devices 3325 are in
communication with the CPU 3305 through a communication bus (solid
lines), such as a motherboard. The storage unit 3315 can be a data
storage unit (or data repository) for storing data.
The CPU 3305 can execute a sequence of machine-readable
instructions, which can be embodied in a program or software. The
instructions may be stored in a memory location, such as the memory
3310. Examples of operations performed by the CPU 3305 can include
fetch, decode, execute, and writeback.
The storage unit 3315 can store files, such as drivers, libraries
and saved programs. The storage unit 3315 can store programs
generated by users and recorded sessions, as well as output(s)
associated with the programs. The storage unit 3315 can store user
data, e.g., user preferences and user programs. The computer system
3301 in some cases can include one or more additional data storage
units that are external to the computer system 3301, such as
located on a remote server that is in communication with the
computer system 3301 through an intranet or the Internet.
The computer system 3301 can be in communication with an OCM and/or
ETL system 3330, including an OCM and/or ETL reactor and various
process elements. Such process elements can include sensors, flow
regulators (e.g., valves), and pumping systems that are configured
to direct a fluid.
Methods as described herein can be implemented by way of machine
(e.g., computer processor) executable code stored on an electronic
storage location of the computer system 3301, such as, for example,
on the memory 3310 or electronic storage unit 3315. The machine
executable or machine readable code can be provided in the form of
software. During use, the code can be executed by the processor
3305. In some cases, the code can be retrieved from the storage
unit 3315 and stored on the memory 3310 for ready access by the
processor 3305. In some situations, the electronic storage unit
3315 can be precluded, and machine-executable instructions are
stored on memory 3310.
The code can be pre-compiled and configured for use with a machine
have a processor adapted to execute the code, or can be compiled
during runtime. The code can be supplied in a programming language
that can be selected to enable the code to execute in a
pre-compiled or as-compiled fashion.
Aspects of the systems and methods provided herein, such as the
computer system 3301, can be embodied in programming. Various
aspects of the technology may be thought of as "products" or
"articles of manufacture" typically in the form of machine (or
processor) executable code and/or associated data that is carried
on or embodied in a type of machine readable medium.
Machine-executable code can be stored on an electronic storage
unit, such memory (e.g., read-only memory, random-access memory,
flash memory) or a hard disk. "Storage" type media can include any
or all of the tangible memory of the computers, processors or the
like, or associated modules thereof, such as various semiconductor
memories, tape drives, disk drives and the like, which may provide
non-transitory storage at any time for the software programming.
All or portions of the software may at times be communicated
through the Internet or various other telecommunication networks.
Such communications, for example, may enable loading of the
software from one computer or processor into another, for example,
from a management server or host computer into the computer
platform of an application server. Thus, another type of media that
may bear the software elements includes optical, electrical and
electromagnetic waves, such as used across physical interfaces
between local devices, through wired and optical landline networks
and over various air-links. The physical elements that carry such
waves, such as wired or wireless links, optical links or the like,
also may be considered as media bearing the software. As used
herein, unless restricted to non-transitory, tangible "storage"
media, terms such as computer or machine "readable medium" refer to
any medium that participates in providing instructions to a
processor for execution.
Hence, a machine readable medium, such as computer-executable code,
may take many forms, including but not limited to, a tangible
storage medium, a carrier wave medium or physical transmission
medium. Non-volatile storage media include, for example, optical or
magnetic disks, such as any of the storage devices in any
computer(s) or the like, such as may be used to implement the
databases, etc. shown in the drawings. Volatile storage media
include dynamic memory, such as main memory of such a computer
platform. Tangible transmission media include coaxial cables;
copper wire and fiber optics, including the wires that comprise a
bus within a computer system. Carrier-wave transmission media may
take the form of electric or electromagnetic signals, or acoustic
or light waves such as those generated during radio frequency (RF)
and infrared (IR) data communications. Common forms of
computer-readable media therefore include for example: a floppy
disk, a flexible disk, hard disk, magnetic tape, any other magnetic
medium, a CD-ROM, DVD or DVD-ROM, any other optical medium, punch
cards paper tape, any other physical storage medium with patterns
of holes, a RAM, a ROM, a PROM and EPROM, a FLASH-EPROM, any other
memory chip or cartridge, a carrier wave transporting data or
instructions, cables or links transporting such a carrier wave, or
any other medium from which a computer may read programming code
and/or data. Many of these forms of computer readable media may be
involved in carrying one or more sequences of one or more
instructions to a processor for execution.
EXAMPLES
Example 1
Fuel Production from OCM Produced Ethylene
An example liquid fuel production process is shown in FIG. 34 and
described in greater detail below. In this example, an OCM product
gas containing ethylene 3402 is preheated to 200.degree. C. to
500.degree. C. depending upon the desired process. The ethylene may
be from 0.05% to 100% pure. For less than 100% pure, the ethylene
containing gas may include CO.sub.2, CO, H.sub.2, H.sub.2O,
C.sub.2H.sub.6, CH.sub.4, C.sub.3 or higher hydrocarbons (i.e.,
C.sub.3+ hydrocarbons), or combinations thereof.
The heated ethylene containing gas 3402 is then flowed through one
or more ethylene conversion reactors, e.g., reactors 3404, 3406 and
3408, each containing a solid acid catalyst. The different reactors
may include reactors having the same catalyst for performing a
parallel reaction to produce a single product. Alternatively, and
in accordance with certain aspects of the invention, the different
reactors may include different catalysts and/or be operated under
different reaction conditions to produce different reaction
products or product ranges. The catalysts may include crystalline
catalysts, such as zeolites, e.g., zeolites ZSM-5, Y, Beta, ZSM-22,
ZSM-48, SAPO-34, SAPO-5, SAPO-11, Mordenite, Ferrierite, and
others. Alternatively or additionally, the catalysts may include
crystalline mesoporous materials, such as SBA-15, SBA-16, MCM-22,
MCM-41, and Al-MCM-41 catalysts, among others. Zeolites and
mesoporous materials can be modified with metals, metal oxides, or
metal ions to enhance ethylene reactivity, product slate
selectivity, and/or catalyst stability.
The ethylene reacts with the solid catalyst to make higher carbon
oligomers/products (C.sub.3-C.sub.30). Carbon number ranges can be
targeted depending on catalyst type and process conditions.
The oligomerized ethylene product stream 3412 exits from the
ethylene conversion reactor(s) and may be used to heat the incoming
ethylene containing gas 3402, e.g., via a heat exchanger 3414. The
product stream is otherwise passed through a series of heat
exchangers 3416, 3418, and 3420 to cool the oligomerized product
and to generate steam 3422. The product stream 3412 is then passed
through a flash drum 3424 to condense heavier products into liquids
3426. Light products 3436, such as C.sub.3-C.sub.4's, can be
recycled back to the ethylene conversion reactor in stream 3428
through compressor 3438 for reaction if the C.sub.3-C.sub.4's are
olefinic and/or to control the heat of reaction of the ethylene
conversion reactors 3404, 3406 and 3408. Alternatively, they may be
routed through downstream processes, e.g., through hydrogenation
reactor 3430 in stream 3436. If desired, the liquid fraction 3426
is passed through a hydrogenation reactor 3430 to hydrogenate
olefins to paraffins/isoparaffins using a Co/Mo, Pd, Ni/Mo or other
hydrogenation catalyst. The oligomerized product 3426 (or
optionally hydrogenated fraction 3432) may then be routed to a
distillation column 3434 to fractionate different cuts of products
3440, such as gasoline, jet, and diesel fuel, fuel blendstocks or
aromatics.
Example 2
Performance of an ETL Reaction
In another example, the performance of an ETL reaction is assessed.
The ETL reaction is performed in an ETL reactor to yield a gasoline
product. Ethylene is introduced to a packed bed of extruded
H--Mg-ZSM-5 catalyst at a WHSV of about 0.7 hr.sup.-1 and a
temperature of about 350.degree. C. The ethylene partial pressure
is about 1 bar. The reactor effluent is chilled to a condenser
temperature of about -5.degree. C., and a portion of the
non-condensing vapors are recycled by a gas pump to the reactor
inlet. In this example, the volumetric recycle-to-feed ratio is
about 2:1. The feed to the reactor is comprised of the combination
of ethylene feed and recycled vapors, yielding an ethylene
concentration at the reactor inlet of about 25%. The product slate
and performance of the catalyst prepared according to the methods
described herein are detailed in FIG. 5.
Example 3
ETL Reaction Products
ETL processes are conducted as described in this disclosure, and
the reaction product properties are measured. During the ETL
oligomerization process, a small amount of coke is produced. Over
time, the coke will deactivate the catalyst below desired levels.
Catalyst activity can be restored to full activity by removing the
coke by oxidation. The catalyst is robust to coke and decoke
cycles. As the catalyst deactivates, the product slate changes. A
freshly regenerated catalyst bed will be more selective to
aromatics and paraffins. Overtime, the catalyst bed will become
less selective toward aromatics and paraffins and more selective
toward olefins. FIG. 35 shows the effect that catalyst time on
stream for a single reactor has on the product slate composition.
Time on stream (TOS) progresses along the x-axis from start of run
(SOR) to end of run (EOR), and the width of product bands on the
y-axis shows their relative abundance. From top to bottom, the
products shown are C.sub.10+ compounds, aromatics, naphthenes,
isoparaffins, N-paraffins, propane/butane, propene/butene, C.sub.5+
olefins, and other compounds.
ETL processes are conducted with different feedstocks, and the
reactor output is compared with PIONA analysis (paraffin content,
isoparaffin content, olefins content, naphthenes content, and
aromatics content), as shown in FIGS. 36A-36E. The feedstocks
compared are ethylene (FIG. 36A), propylene (FIG. 36B), butylene
(FIG. 36C), 50:50 ethylene/propylene (FIG. 36D), and 50:50
ethylene/butylene (FIG. 36E). The liquid products from the
different feeds are comparable in composition and carbon number
distribution, showing the robustness of the process with respect to
feed composition.
ETL processes are conducted at different peak catalyst bed
temperatures, and the effect on product composition is evaluated,
as shown in FIG. 37. The x-axis shows the temperature from
315.degree. C. to 385.degree. C., and the y-axis shows the liquid
mol % of various product components. From top to bottom, the
product components are C10+ compounds, aromatics, naphthenes,
olefin, isoparaffins, and paraffins.
Different segments of ETL product components can be directed for
use in different fractions. For example, a separations process can
be employed to separate a jet fraction (comprising, e.g., C.sub.10+
compounds) from a gasoline fraction (comprising, e.g., C.sub.9-
compounds), as shown in FIG. 38. In some cases, the ETL product
stream can comprise about 65% gasoline fraction components and
about 35% jet fraction components.
Reactor operating conditions can impact the reactor performance,
and can favor the production of components for a particular product
slate. For example, operating conditions and reactor performance
for the production can be those shown in Table 3, favoring the
production of gasoline components. The resulting product can have a
stream composition as shown in Table 4, and can be characterized by
the properties shown in Table 5 (center column), with reference to
the specification for RBOB (left column).
In another example, operating conditions can favor the production
of aromatics, such as the operating conditions and reactor
performance shown in Table 6. The resulting product can have a
stream composition as shown in Table 7.
TABLE-US-00003 TABLE 3 Operating conditions and reactor
performance, gasoline ETL Gasoline Inlet T (.degree. C.) 300 Outlet
T (.degree. C.) 383 Inlet P (Barg) 25 Outlet P (Barg) 25 WHSV
(h.sup.-1) 1.4 C.sub.2= conversion >99% C.sub.5+ Selectivity 63%
Composition mol % Inlet CH.sub.4 95 C.sub.2H.sub.4 5 C balance
99.1%
TABLE-US-00004 TABLE 4 Outlet stream composition, gasoline (in C
mol %) n-paraffins i-paraffins olefins napthenes aromatics Total C1
0.00% 0.00% 0.00% 0.00% 0.00% 0.00% C2 0.00% 0.00% 0.00% 0.00%
0.00% 0.00% C3 9.62% 0.00% 1.90% 0.00% 0.00% 11.52% C4 7.51% 14.92%
2.93% 0.00% 0.00% 25.36% C5 2.87% 8.20% 1.70% 2.00% 0.00% 14.77% C6
0.45% 6.88% 1.57% 1.46% 0.00% 10.36% C7 0.16% 3.34% 0.91% 0.90%
3.35% 8.65% C8 0.08% 1.02% 0.26% 1.06% 6.63% 9.06% C9 0.03% 0.91%
0.24% 0.48% 6.53% 8.19% C10 0.03% 0.76% 0.02% 0.18% 4.14% 5.12%
C11+ 6.96%
TABLE-US-00005 TABLE 5 Gasoline fuel properties RBOB product
specification properties Chemical [max.] Benzene (Vol %) 1.30%
0.97% Aromatics (Vol %) 50% 35.31% Olefins (Vol %) 25% 24.6% Octane
[min] RON -- 96.9 MON 82 84.9 Tot. Octane 87 90.9 Distillation
[max] RVP (psi) 15 9.37 10% (.degree. C.) 70 57.67 50% (.degree.
C.) 121 113.78 90% (.degree. C.) 190 161.28 FBP (.degree. C.) 221
192.39 Oxidation stability Induction time (min) 240 >240
TABLE-US-00006 TABLE 6 Operating conditions and reactor
performance, aromatics ETL Gasoline Inlet T (.degree. C.) 300
Outlet T (.degree. C.) 383 Inlet P (Barg) 25 Outlet P (Barg) 25
WHSV (h.sup.-1) 1.4 C.sub.2= conversion >99% C.sub.5+
Selectivity 63% Composition mol % Inlet CH.sub.4 95 C.sub.2H.sub.4
5 C balance 99.1%
TABLE-US-00007 TABLE 7 Outlet stream composition, aromatics (in C
mol %) n-paraffins i-paraffins olefins napthenes aromatics Total C1
0.00% 0.00% 0.00% 0.00% 0.00% 0.00% C2 0.00% 0.00% 0.00% 0.00%
0.00% 0.00% C3 9.62% 0.00% 1.90% 0.00% 0.00% 11.52% C4 7.51% 14.92%
2.93% 0.00% 0.00% 25.36% C5 2.87% 8.20% 1.70% 2.00% 0.00% 14.77% C6
0.45% 6.88% 1.57% 1.46% 0.00% 10.36% C7 0.16% 3.34% 0.91% 0.90%
3.35% 8.65% C8 0.08% 1.02% 0.26% 1.06% 6.63% 9.06% C9 0.03% 0.91%
0.24% 0.48% 6.53% 8.19% C10 0.03% 0.76% 0.02% 0.18% 4.14% 5.12%
C11+ 6.96%
Example 4
ETL Catalyst Formation and Use
To form ETL catalyst, base material, dopant, and binder are mixed
in desired ratios and then extruded. The target catalyst form
strength is to maintain particle crush strength above 3 N/mm. The
crush strength threshold is selected based upon the expected stress
on individual particles in a commercial scale reactor. As shown in
FIG. 39, the ETL catalyst baseline formulation is above the crush
strength threshold. If desired, stronger catalyst forms can be
achieved by tailoring the active catalyst to binder ratio.
Catalyst aging can be accelerated by changing process conditions,
such as WHSV. Catalyst aging can be accelerated without changing
process inputs. FIG. 40 shows catalyst aging measured by fractional
ethylene conversion (y-axis) as a function of time on stream (TOS,
x-axis) in hours. Catalyst aging under typical commercial
conditions is shown by the curve on the right, and catalyst aging
under accelerated conditions is shown by the curve on the left.
The catalyst used can be robust to a number of process and
regeneration cycles, with little to no impact on the product
composition. FIG. 41 shows the product composition produced by a
reactor in its first cycle, i.e. no catalyst regeneration (left
side) compared to a reactor in its tenth cycle, i.e. nine catalyst
regenerations (right side). The product components graphed, from
top to bottom, are C.sub.11+ compounds, aromatics, naphthenes,
olefins, iso-paraffins, and paraffins.
Systems and methods of the present disclosure can be combined with
or modified by other systems and methods, such as those described
in U.S. Pat. No. 2,943,125, U.S. Pat. No. 3,686,350, U.S. Pat. No.
4,101,600, U.S. Pat. No. 8,624,042, and U.S. Pat. No. 5,792,895,
U.S. patent application Ser. No. 14/099,614 and PCT/US2013/073657,
each of which is entirely incorporated herein by reference.
It should be understood from the foregoing that, while particular
implementations have been illustrated and described, various
modifications can be made thereto and are contemplated herein. It
is also not intended that the invention be limited by the specific
examples provided within the specification. While the invention has
been described with reference to the aforementioned specification,
the descriptions and illustrations of the preferable embodiments
herein are not meant to be construed in a limiting sense.
Furthermore, it shall be understood that all aspects of the
invention are not limited to the specific depictions,
configurations or relative proportions set forth herein which
depend upon a variety of conditions and variables. Various
modifications in form and detail of the embodiments of the
invention will be apparent to a person skilled in the art. It is
therefore contemplated that the invention shall also cover any such
modifications, variations and equivalents. It is intended that the
following claims define the scope of the invention and that methods
and structures within the scope of these claims and their
equivalents be covered thereby.
* * * * *